Integrated reactors, methods of making same, and methods of conducting simultaneous exothermic and endothermic reactions

ABSTRACT

Integrated Combustion Reactors (ICRs) and methods of making ICRs are described in which combustion chambers (or channels) are in direct thermal contact to reaction chambers for an endothermic reaction. Superior results were achieved for combustion chambers which contained a gap for flee flow through the chamber. Particular reactor designs are also described. Processes of conducting reactions in integrated combustion reactors are described and results presented. Some of these processes are characterized by unexpected and superior results.

RELATED APPLICATIONS

This application is a divisional of U.S. patent application Ser. No.10/076,875, now U.S. Pat. No., which was a continuation-in-part of U.S.patent application Ser. No. 09/375,614, now U.S. Pat. No. 6,488,838 andSer. No. 09/640,903, now U.S. Pat. No. 6,680,044, which are incorporatedherein as if reproduced in full below. In accordance with 35 U.S.C.sect. 119(e), this application claims priority to U.S. ProvisionalApplication No. 60/269,628, filed Feb. 16, 2001.

FIELD OF THE INVENTION

The present invention relates to integrated reactors for conductingexothermic and endothermic reactions, methods of making integratedreactors, and methods of conducting reactions in integrated reactors.

Introduction

Currently, endothermic reactions performed in microreactors are drivenusing heat from an external source, such as the effluent from anexternal combustor. In doing so, the temperature of the gas streamproviding the heat is limited by constraints imposed by the materials ofconstruction. For example, a typical microreactor constructed fromInconel 625 might be limited in use for gas service to temperatures of˜1050° C. or less. Practically, this means that tile effluent from anexternal combustor must be diluted with cool gas (i.e. excess air) tobring the gas temperature down to meet material temperature constraints.This increases the total gas flow rate, raising blower/compressor costs.Moreover, heating the gas stream externally introduces heat losses(associated with delivery of the hot gas to the microreactor) andexpensive high temperature materials to connect the combustor to themicroreactor.

On the other hand, an integrated combustor can produce heat for thereaction in close proximity to the reaction zone, thus reducing heatlosses and increasing efficiency. Because traditional combustioncatalysts are riot stable at high temperatures (above ˜1200° C.) due tonoble metal sintering, the integrated combustor must remove heat at arate sufficient to keep local temperatures at the catalyst surface belowthis level or risk rapid catalyst deactivation.

SUMMARY OF THE INVENTION

In an integrated reactor, combustion/heat generation should occur inclose proximity to the endothermic reaction. Preferably, an exothermicreaction occurs in microchannels that are interleaved with microchannelsin which there is an endothermic reaction. Co-flow of endothermic andexothermic reaction streams is preferred; however, cross-flow orcountercurrent flow is also an option. The heat of an exothermicreaction is conducted from the exothermic reaction catalyst to theendothermic reaction catalyst, where it drives the endothermic reaction.This rapid heat removal from the combustion region enables the option touse a very small fraction of excess air (e.g., close to stoichiometriccombustion, which could produce temperatures exceeding 1400° C. ifreacted adiabatically). The use of a flow-by catalyst configuration forone or both the exothermic and endothermic microchannels can create anadvantageous capacity/pressure drop relationship. In a flow-by catalystconfiguration, gas flows in a 0.05-2 mm gap adjacent to a thin layer ofengineered catalyst that contacts the microchannel wall. The catalystmay be either inserted and adjacent to the reactor wall or integral withthe reactor wall. In the case of integral with the reactor wall, apreferred method is washcoating a catalyst on a wall or walls of themicrochannel. The catalyst may include the use of additional layers forincreasing surface area, such as a porous high surface area ceramic, orlayers for promoting adhesion of a ceramic to metal, such as amorphoustitania that is either CVD or solution deposited. The use of channelshaving a minimum dimension of more than 2 mm may be less effective sinceheat and mass transfer limitations may be magnified. An integratedcombustor can use the high surface area of microreactor channels toremove heat as it is produced, thus keeping microreactor components fromexceeding material temperature constraints while combusting with muchless excess air (or diluent) than would be necessary for an externalcombustor.

In one aspect, the invention provides an integrated reactor, thatincludes: a first reaction chamber having a width of 2 mm or less, wherethere is an open channel through the first reaction chamber, wherein thefirst reaction chamber has an internal volume comprising 5 to 95 vol. %of porous catalyst and 5 to 95 vol. % of open space. The integratedreactor also includes a second reaction chamber having a width of 2 mmor less, wherein there is an open channel through the second reactionchamber, wherein the second reaction chamber has an internal volumecomprising a catalyst and at least 5 vol. % of open space; and areaction chamber wall separating the first chamber and the secondchamber. This integrated reactor possesses a heat flux characteristic ofat least 1 W/cc as measured according to the Heat Flux Measurement Test.

Tile invention also includes methods of performing exothermic andendothermic reactions in the reactor. An exothermic reaction compositionis a chemical composition that will react under the selected conditionsto produce heat; typically a catalyst will catalyze the reaction.

In another aspect, the invention provides an integrated reactor, thatincludes: a first reaction chamber having a width of 2 mm or less,wherein there is an open channel through the first reaction chamber,wherein the first reaction chamber has an internal volume comprising 5to 95 vol. % of porous catalyst and 5 to 95 vol. % of open space; and asecond reaction chamber having a width of 2 mm or less, wherein there isan open channel through the second reaction chamber, wherein the secondreaction chamber has an internal volume comprising a catalyst and atleast 5 vol. % of open space. A reaction chamber wall separates thefirst chamber and the second chamber; and the integrated reactorpossesses a NO_(x) output characteristic of less than 100 ppm asmeasured according to the Standard NO_(x) Test Measurement.

The invention also provides a method of making an integrated reactor,comprising: providing a single block of thermally conductive material;forming at least one first microchannel in the block; forming at leastone second microchannel in the block; placing at least one catalystcapable of catalyzing an exothermic reaction in the at least one firstmicrochannel; and placing at least one catalyst capable of catalyzing anendothermic reaction in the at least one second microchannel. In theintegrated reactor, the first microchannel and second microchannel areseparated by less than 1 cm.

The invention further provides a method of conducting an endothermicreaction in an integrated combustion reaction, comprising: passing anexothermically reacting composition into at least one exothermicreaction chamber, wherein the exothermic reaction chamber comprises atleast one exothermic reaction chamber wall that is adjacent at least oneendothermic reaction chamber, wherein the combustion chamber comprises aexothermic reaction catalyst in contact with at least the at least oneexothermic reaction chamber wall that is adjacent at least oneendothermic reaction chamber, wherein the exothermic reaction catalysthas an exposed surface within the exothermic reaction chamber, andwherein the exposed surface of the exothermic reaction catalyst and asecond surface within the exothermic reaction chamber define an openchannel within the exothermic reaction chamber, wherein the gap has athickness, in a direction perpendicular to net flow where the directionof net flow is tile direction that gas would travel through thecombustion chamber during operation, of 2 mm or less; wherein theexothermic reaction composition reacts in the exothermic reactionchamber and generates heat; and passing an endothermic reaction mixtureinto the at least one endothermic reaction chamber; and wherein themethod has a volumetric heat flux of at least 1 W/cc.

In another aspect, the invention provides a method of conducting anendothermic reaction in an integrated combustion reaction, comprising:passing an endothermic reaction composition into at least oneendothermic reaction chamber, passing an exothermic reaction compositioninto at least one exothermic reaction chamber, wherein the exothermicreaction chamber comprises at least one exothermic reaction chamber wallthat is adjacent at least one endothermic reaction chamber, wherein theendothermic reaction chamber comprises an endothermic reaction catalystin contact with at least the at least one endothermic reaction chamberwall that is adjacent at least one exothermic reaction chamber, whereintile endothermic reaction catalyst comprises an exposed surface withinthe endothermic reaction chamber, and wherein the exposed surface of theendothermic reaction catalyst and a second surface within theendothermic reaction chamber define a gap within the endothermicreaction chamber, wherein the gap has a thickness, in a directionperpendicular to net flow where the direction of net flow is thedirection that gas would travel through the endothermic chamber duringoperation, of 2 mm or less; and wherein the method is controlled suchthat heat flux between the at least one exothermic chamber and the atleast one endothermic reaction chamber is 1 W/cc or more.

The invention also provides a method of conducting an endothermicreaction in an integrated combustion reaction, comprising: passing anendothermic reaction composition into at least one endothermic reactionchamber, passing an exothermic reaction composition into at least oneexothermic reaction chamber, wherein the exothermic reaction chambercomprises at least one exothermic reaction chamber wall that is adjacentat least one endothermic reaction chamber, wherein the endothermicreaction chamber comprises an endothermic reaction catalyst in contactwith at least the at least one endothermic reaction chamber wall that isadjacent at least one exothermic reaction chamber, wherein theendothermic reaction catalyst comprises an exposed surface within theendothermic reaction chamber, and wherein the exposed surface of theendothermic reaction catalyst and a second surface within theendothermic reaction chamber define a gap within the endothermicreaction chamber, wherein the gap has a thickness, in a directionperpendicular to net flow where the direction of net flow is thedirection that gas would travel through tile endothermic chamber duringoperation, of 2 mm or less; wherein the exothermic reaction compositioncomprises air and a fuel; and wherein the exothermic reactioncomposition is converted to products and the products have less than 100ppm NO_(x).

In another aspect (or in combination with any of the foregoing aspects),the present invention provides an integrated reactor including: at leastone endothermic reaction chamber and/or at least one exothermic chamber,wherein at least one reaction chamber comprises at least one porouscatalyst material and at least one open channel wherein each of the atleast one (exothermic or endothermic) reaction chambers has an internalvolume defined by reaction chamber walls in the direction of height andwidth, and by length of catalyst in tile length direction. The internalvolume has dimensions of chamber height, chamber width and chamberlength. At least one exothermic reaction chamber and at least oneendothermic reaction chamber (which is adjacent the exothermic reactionchamber) comprises a chamber height or chamber width that is about 2 mmor less. At a point where the chamber height or the chamber width isabout 2 mm or less, the chamber height and the chamber width define across-sectional area. The cross-sectional area of at least one reactionchamber comprises a porous catalyst material and an open area, where theporous catalyst material occupies 5% to 95% of thie cross-sectional areaand where the open area occupies 5% to 95% of the cross-sectional area.The open area in the cross-sectional area occupies a contiguous area of5×10⁻⁸ to 1×10⁻² m² and the porous catalyst material has a pore volumeof 5 to 98% and more than 20% of the pore volume comprises pores havingsizes of from 0.1 to 300 microns.

The invention also includes devices having any of the unique structuralfeatures or designs described herein. For example, the inventionincludes apparatus that includes a fuel-air mixing manifold as shown inFIG. 19. The invention also includes processes using any of thestructural features or designs, or characterized by any of theproperties or results described herein.

Various embodiments of the present invention may possess advantages suchas low pressure drop, low requirement for excess air, high combustionstability, short contact time, low CO/NOx formation, operation at nearstiochiometric air feed, greater safety, and high thermal cyclingdurability. Operation with a near stoichiometric air feed reduces theoverall load on the systems air blower or compressor which will lead tosignificant cost savings.

An additional advantage by reducing the combustion temperature (ortemperature of the exothermic reaction) required to drive theendothermic reaction is use of alternate metals or metallurgy such thatlower cost materials or longer device life may be achieved.

Although the combustion may have both homogeneous and heterogeneouscontributions, catalytic combustion in a microchannel (or channel with aminimum open dimension less than the quench diameter) will reduce thecontribution of homogeneous reactions and favor heterogeneous(catalytic) combustion at the wall. This will also further enhancesafety by inhibiting tile gas phase reactions that might otherwise takethe combustion mixture well above the safe operating temperature limitof the material. Inhibition grows stronger with decreasing channelminimum dimension and with increasing catalytic surface area on thechannel walls.

In conjunction with other features of the invention, the use of aflow-by configuration in which a reaction chamber has a gap such thatgases can flow by (rather than through) a catalyst allows for asignificant improvement in performance over the prior art. This improvedperformance was evidenced in the integrated combustion reactor (ICR)tests by a much higher heat flux (e.g., 29 W/cm² on an area basis or 118W/cm³ on a volumetric basis) than any reported in the literature for aminimal pressure drop (e.g., <4 psi (0.3 bar) in a 1 inch reactorlength. Because the ICR can achieve such high heat fluxes withoutcausing excessive pressure drop, endothermic reaction contact times arefeasible which are much shorter than those for flow-through catalyticdevices or monolith devices. Shorter contact times enable a higherproductivity or throughput through unit volume of reactor.

Introduction of laterally distributed (across a channel) combustion fueland air in co-flow with endothermic reactant flow concentrates the heattransfer at the endothermic reactor inlet, where the concentrationgradient (and therefore rate of reaction) is highest; thus obtainingsuperior results over systems that distribute the combustion fuel evenlyover the entire surface of the combustion catalyst. Although theexamples with distributed combustion still exhibit excellent heat fluxin comparison to conventional steam reformers.

It is also recognized that the present invention could use alternateexothermic reactions, such as oxidation reactions, including partialoxidation reaction, to drive an endothermic reaction.

Glossary

“Shims” refer to substantially planar plates or sheets that can have anywidth and height and preferably have a thickness (the smallestdimension) of 2 millimeter (mm) or less, and in some preferredembodiments between 50 and 500 μm.

“Unit operation” means chemical reaction, vaporization, compression,chemical separation, distillation, condensation, heating, or cooling.“Unit operation” does not mean merely mixing or fluid transport,although mixing and transport frequently occur along with unitoperations. A microchannel has at least one dimension of 2 mm or less.

An “open channel” is a gap of at least 0.05 mm that extends all the waythrough a reaction chamber such that gases can flow through the reactionchamber with relatively low pressure drop.

During operation, a reactant enters a combustion or reaction chamber ina bulk flow path flowing past and in contact with a “porous material” or“porous catalyst.” A portion of the reactant molecularly transverselydiffuses into the porous catalyst and reacts to form a product orproducts, and then the product(s) diffuses transversely into the bulkflow path and out of the reactor.

Tile term “bulk flow region” refers to open areas or open channelswithin the reaction chamber. A contiguous bulk flow region allows rapidgas flow through the reaction chamber without large pressure drops. Inpreferred embodiments there is laminar flow in the bulk flow region.Bulk flow regions within each reaction chamber preferably have across-sectional area of 5×10⁻⁸ to 1×10⁻² m², more preferably 5×10⁻⁷ to1×10⁻⁴ m². The bulk flow regions preferably comprise at least 5%, morepreferably 30-80% of either 1) the internal volume of the reactionchamber, or 2) the cross-section of the reaction chamber.

“Equilibrium conversion” is defined in the classical manner, where themaximum attainable conversion is a function of the reactor temperature,pressure, and feed composition. For the case of hydrocarbon steamreforming reactions, the equilibrium conversion increases withincreasing temperature and decreases with increasing pressure.

BRIEF DESCRIPTION OF THE FIGURES

FIG. 1 illustrates an integrated reactor.

FIGS. 2 a and 2 b illustrate reaction chamber designs.

FIGS. 3 and 4 a illustrate cross-sections of integrated reactors.

FIG. 4 b is an exploded view of an integrated reactor.

FIGS. 5 and 6 illustrate integrated steam reforming reactors.

FIG. 7 illustrates shim designs for making a reactor.

FIG. 8 is a schematic illustration of a reactor used in tie Examples.

FIGS. 9-12 are data from the Examples.

FIG. 13 is a schematic illustration of a reactor used in the Examples.

FIGS. 14-16 are data from the Examples.

FIGS. 17-20 illustrate reactor designs of reactors described in theExamples.

FIGS. 21-24 are data from the Examples.

FIG. 25 illustrates a reactor design of a reactor described in theExamples.

FIGS. 26-27 are data from the Examples.

DETAILED DESCRIPTION OF THE INVENTION

An integrated reactor according to the present invention includes afirst reaction chamber that contains a catalyst capable of catalyzing anexothermic reaction and an adjacent second reaction chamber thatcontains a catalyst capable of catalyzing an endothermic reaction. Areaction chamber wall separates the first and second reaction chambers.

An illustration of a portion of a preferred type of reactor apparatus isshown in FIG. 1. Exothermic reaction chamber 12 contains exothermicreaction catalysts 14, 16 and open channel 18. Endothermic reactionchamber 15 includes endothermic reaction catalyst 17 and open channel19.

In the present invention, the exothermic (and/or endothermic) reactionchamber(s) preferably has a width (the smallest dimension that isperpendicular to flow) of 2 mm or less, more preferably 1 mm or less andin some embodiments 0.5 mm or less. The dimensions of a reaction chamberare the internal dimensions and include catalyst but do not includechamber walls. A reaction chamber wall (separating the exothermic andendothermic reaction chambers) should be thermally conductive andpreferably has a thickness (the distance between reaction chambers) of 5mm or less, more preferably 2 mm or less, and in some embodiments awidth of 1 to 2 mm. A short heat transport distance desired for goodperformance. It has been discovered that these short heat transportdistances, combined with preferred reactor configurations, can providesurprisingly high volumetric productivity and low pressure drop.

As described in greater detail below, preferred processes of theinvention can be described by the configuration of the apparatus and/orby measurable characteristics such as heat flux, volumetricproductivity, and/or pressure drop (which could also be described inconjunction with process conditions such as flow rate, temperature,etc.).

FIG. 1 illustrates reaction chambers having the shape ofparallelepipeds; however, it should be appreciated that other shapessuch as cylinders (for example, adjacent cylinders or cylinders with anexothermic catalyst partly surrounded by an arc containing anendothermic reaction catalyst, or vice versa), or prisms (preferablyclose packed prisms to reduce heat transport distance and maximizesurface area for thermal transport). Such shapes could be made, forexample, by drilling through a block or laminating a stack of shims withshapes, aligned apertures such that the stacked and bonded shims form apassage having borders defined by the edges of the apertures. Toincrease surface area, in some embodiments, the reaction chamber (eitherexothermic, endothermic, or both) can have a projections or a set ofmicrochannels such as shown in FIGS. 2 a and 2 b. The projections 22 ormicrochannel walls 24 can be coated with a catalyst to form a catalystcoating (not shown) such as by wash coating, chemical vapor coating,etc.

In another preferred embodiment (see FIG. 3), a reaction chamber wall 31has fins 32. The fins can have any shape and can extend partly orcompletely over the width of a reaction chamber. Preferably, a catalystor catalysts (not shown) are deposited over the reaction chamber wallsto form exothermic or endothermic reaction chambers. In the illustrateddevice, flow during operation is into or out of the page. Preferably thereaction chambers are stacked in alternating layers or exothermic 34 andendothermic 36 reaction chambers separated by a thermally conductivechamber wall 38.

Alternating layers of exothermic and endothermic reaction chambers is ageneral feature of the invention, and preferably there are at least 2,more preferably at least 5 layers of endothermic reaction chambersalternating with at least 1, more preferably at least 4 layers ofexothermic reaction chambers. Preferably, the apparatus is designed, andthe methods performed such that outer layers have one half the flow ofreactants as compared with inner layers of the same type; for example,in a device having 2 exothermic reaction chambers interleaved between 3endothermic reaction chambers, the outer 2 endothermic reaction chambershave one half the flow of the inner endothermic reaction chamber.

In another embodiment, see FIG. 4 a, and integrated reaction chamber isformed by adjacent reaction chambers in which one reaction chambercontains a porous catalyst insert 42 (in FIG. 4 a, the insert fills thereaction chamber such that reactants would flow through the catalyst—aless preferred embodiment that can be utilized where relatively largepressure drops are acceptable or the catalyst reaction kinetics do notsupport very fast reactions of lens of milliseconds). Reactor walls 43,45 separate chamber 42 from the adjacent reaction chamber 44 that hasopen channels that may have a catalyst (not shown) disposed in eachchannel. FIG. 4 b shows an exploded view with the pieces that can beassembled into the bonded device.

FIG. 5 shows an embodiment which could, for example, be for a steamreformer in which hydrogen (or alternate fuel) is combusted to supplyheat. A reformate composition flows through open channel 52. Hydrogen(or alternate combustion fuel) flows through open channel 54. Air (orother oxygen-containing composition) in channels 56 flows throughaperture 58 and into channel 54 where the oxygen reacts with hydrogen,chiefly on combustion catalyst 53. Preferably, all dimensions in thehydrogen channel are less than or near the quench distance for hydrogen(or alternate fuel, which could comprise hydrogen and otherhydrocarbons) and the selected oxidant at the required temperature,although this may not be possible especially in the case of a hydrogenfuel. Combustion exhaust 55 preheats the oxygen-containing composition.Heat of combustion is supplied to the reforming catalyst 57 throughreaction chamber wall 59. Desirably, the dimensions of open channel 54are controlled to suppress flame formation. In this embodiment, theoxygen-containing composition is introduced in the front (i.e., near thehydrogen inlet) of the reactor and the combustion gas and reformate flowin a coflow arrangement. This configuration has two significantadvantages: (1) both the exothermic and endothermic reactors have thehighest concentration of reactants and therefore fastest reaction in thesame area, and (2) the hydrogen is mixed with oxygen within the reactionchamber, thus avoiding premixing and reaction outside the reactionchamber.

FIG. 6 shows a similar arrangement except that the wall separating thehydrogen channel and the air or oxygen channel has multiple apertures 62that distribute oxygen throughout the hydrogen channel. FIG. 7 is anoverhead view of plates 64, 66 that could be used to construct thehydrogen channel(s) 61 and separator plate 64 (either plate could beused). In a variation of the illustrated process in FIG. 6, hydrogencould be distributed into an air (or oxygen) chamber that is adjacentthe endothermic reaction chamber.

Of course, it should be appreciated that in any of the devices describedherein, alternative reactants could be used in place of any of thereactants mentioned. For example, methane or other fuel could be used inplace of hydrogen. Similarly, conditions can be varied, for example,flow could be cross or counter-flow. Although, in some figures,apparatus is shown with only one exothermic reaction chamber, it shouldbe appreciated that multiple alternating layers could be used, or twoexothermic reaction chambers could sandwich an endothermic reactionchamber.

A reaction chamber has dimensions of height, width and length. Theheight and/or width is preferably about 2 mm or less, and morepreferably 1 mm or less (in which case the reaction chamber falls withinthe classical definition of a microchannel). The length of the reactionchamber is typically longer. Preferably, the length of the reactionchamber is greater than 1 cm, more preferably in the range of 1 to 50cm. Typically, the sides of the reaction chamber are defined by reactionchamber walls. These walls are preferably made of a hard material suchas a ceramic, an iron based alloy such as steel, or monel, or hightemperature nickel based superalloys such as inconel 625, inconel 617 orHaynes 230. More preferably, the reaction chamber walls are comprised ofstainless steel or inconel which is durable and has good thermalconductivity.

In addition to thermal transfer between adjacent reaction chambers, insome embodiments, a reaction chamber can be in thermal contact with amicrochannel heat exchanger. This combination of reaction chamber(s) andheat exchanger(s) can result in high rates of thermal transfer. Examplesand more detailed description including the use of microchannel heatexchangers are provided in U.S. patent application Ser. No. 09/492,246,filed Jan. 27, 2000, now U.S. Pat. No. 6,616,909 incorporated herein byreference. In some embodiments, the reaction chamber(s) and heatexchangers have a heat flux of at least 0.6 W per cubic centimeter ofreactor volume.

In some preferred embodiments, the reaction chamber has an inlet and anoutlet with a contiguous bulk flow path from the inlet to the outlet. Inthese preferred embodiments, the pressure drop from inlet to outlet ispreferably less than 20%, more preferably less than 10% of system inletpressure. The pressure drop is preferably less than 350 kPa, and morepreferably the pressure drop is less than 70 kPa. A low pressure drop isdesired to reduce tile size and cost of other system equipment such aspumps and compressors. In other less preferred embodiments, the reactionchamber may include a section, such as a porous plug, that interfereswith bulk flow.

The integrated reactor works best with specific header designs that 1)prevent combustion reaction upstream of the microchannel catalyst, and2) uniformly distribute one of the combustion reactants across tilemicrochannel cross-section.

Preferably, the width of the bulk flow path (open channel gap) within areaction chamber is less than or equal to 1 mm and the length (directionof net flow) is preferably less than or equal to 20 inches (50 cm). Thewidths of the porous catalyst may vary but is preferably at least 20%and more preferably 50% of the circumference of the bulk flow path.

The present invention could also be used for liquid phase reactions. Inthe case of liquid phase reactions, the critical channel dimension willlikely be smaller than that for gas phase reactions to accommodate thereduced mass diffusion rate that brings reactants to the catalyticsurface.

The “porous catalyst material” described herein refers to a porousmaterial having, a pore volume of 5 to 98%, more preferably 30 to 95% oftile total porous mnaterial's volume. At least 20% (more preferably atleast 50%) of the material's pore volume is composed of pores in thesize (diameter) range of 0.1 to 300 microns, more preferably 0.3 to 200microns, and still more preferably 1 to 100 microns. Pore volume andpole size distribution are measured by Mercury porisimetry (assumingcylindrical geometry of the pores) and nitrogen adsorption. As is known,mercury porisimetry and nitrogen adsorption are complementary techniqueswith mercury porisimetry being more accurate for measuring large poresizes (larger than 30 nm) and nitrogen adsorption more accurate forsmall pores (less than 50 nm). Pore sizes in the range of about 0.1 to300 microns enable molecules to diffuse molecularly through thematerials under most gas phase catalysis conditions. The porous materialcan itself be a catalyst, but more preferably the porous materialcomprises a metal, ceramic or composite support having a layer or layersof a catalyst material or materials deposited thereon. The porosity canbe geometrically regular as in a honeycomb or parallel pore structure,or porosity may be geometrically tortuous or random. Preferably thesupport is a foam metal, foam ceramic, metal felt (i.e., matted,nonwoven fibers), or metal screen.

Preferred major active constituents of the catalysts include: elementsin the IUPAC Group IIA, IVA, VA, VIA, VIIA, VIIA, IB, IIB, IVB,Lanthanide series and Actinide series. The catalyst layers, if present,are preferably also porous. The average pore size (volume average) ofthe catalyst layer(s) is preferably smaller than the average pore sizeof the support. The average pore sizes in the catalyst layer(s) disposedupon the support preferably ranges from 10⁻⁹ m to 10⁻⁷ m as measured byN₂ adsorption with BET method. More preferably, at least 50 volume % ofthe total pore volume is composed of pores in the size range of 10⁻⁹ mto 10⁻⁷ m in diameter. Diffusion within these small pores in thecatalyst layer(s) is typically Knudsen in nature for gas phase systems,whereby the molecules collide with the walls of tile pores morefrequently than with other gas phase molecules.

In preferred embodiments, catalysts are in the form of inserts that canbe conveniently inserted and removed from a reaction chamber. Reactionchambers (either of the same type or of different types) can be combinedin series with multiple types of catalysts. For example, reactants canbe passed through a first reaction chamber containing a first type ofcatalyst, and the products from this chamber passed into a subsequentreaction chamber (or a subsequent stage of the same reaction chamber)containing a second type of catalyst in which the product (or morecorrectly termed, the intermediate) is converted to a more desiredproduct. If desired, additional reactant(s) can be added to thesubsequent reaction chamber.

The catalyst could also be applied by other methods such as washcoating. On metal surfaces, it is preferred to first apply a bufferlayer by chemical vapor deposition, thermal oxidation, etc. whichimproves adhesion of subsequent wash coats.

Preferred reactors and methods of conducting reactions in integratedreactors can be characterized by their properties. Unless specifiedotherwise, these properties are measured using the testing conditionsdescribed in the Examples section. The invention can be characterized byany of the properties individually or in any combination. Averagevolumetric heat flux is preferably at least 1 W/cc, or, in otherpreferred embodiments, at least 5, or 10, or 20, or 50, or 100, or about120 W/cc, and in some embodiments between 10 and about 120 W/cc. Thedevices can be characterized by the low NO_(x) output when measured bythe standard NO_(x) test measurement that is described in the Examplessection. NO_(x) output is preferably less than 100 ppm, more preferablyless than 50 ppm, still more preferably less than 10 ppm, and still morepreferably less than 5 ppm, and in some embodiments, NO_(x) output is inthe range of about 5 to 20 ppm. The inventive processes involvingcombustion preferably use less than 100% excess air (or, equivalently,excess oxygen), more preferably less than 75%, still more preferablyless than 50%, yet still more preferably less than 25%, or 10% or 5%excess air. For characterizing devices, excess oxygen is measured underthe conditions set forth in the Heat Flux Measurement Test or (ifcharactrerized in conjuction with NO_(x) output) measured under theconditions set forth in the standard NO_(x) test measurement. Pressuredrop through the exothermic and/or endothermic reaction chambers ispreferably less than the following (in order of preference, based onlength of reaction chamber) 295,000; 250,000; 125,000; 50,000; 25,000;12,500; 2500; or 1500 Pa/cm. For devices, the pressure drop is measuredunder the conditions set forth in the Heat Flux Measurement Test or (ifcharactrerized in conjunction with NO_(x) output) measured under theconditions set forth in the standard NO_(x) test measurement.

The devices may be made of materials such as plastic, metal, ceramic andcomposites, depending on the desired characteristics. Walls separatingthe device from the environment may be thermally insulating; however,the walls separating adjacent exothermic and endothermic reactionchambers should be thermally conductive.

The devices can be made by forming chambers within a single block ofmaterial, by joining multiple components, and by stacking and bondingshims.

A preferred integrated reactor body can be made from a single block ofmetal. Its chapels could be created with a wire EDM in the main body,and the headers and footers could be made separately and welded on,adding to tie flexibility of die design. Wire EDM is used to createslots or holes in a block of metal that are the microchannels throughwhich flow passes and a unit operation occurs. Sinker EDM, lasermachining, and in some larger channels conventional milling can also beused to make channels from a single block of metal.

The aperture-containing shims can be formed by processes including:conventional machining, wire EDM, laser cutting, photochemicalmachining, electrochemical machining, molding, water jet, stamping,etching (for example, chemical, photochemical and plasma etch) andcombinations thereof. For low cost, stamping is especially desirable.The shims may be joined together by diffusion bonding methods such as aram press or a HIP chamber. They may also be joined together by reactivemetal bonding or other methods that create a face seal. Alternately,laser welding shims could join the devices or sheets to form sealsbetween flow paths. Devices could alternatively be joined by the use ofadhesives. In preferred embodiments, devices are laminated in a singlestep, in less preferred embodiments, a first set of shims is bondedtogether and subsequently bonded to a second (or more) set of shims. Insome preferred embodiments, a set of shims is bonded together in asingle step and then the resulting bonded article is cut into multipledevices.

Catalytic processes of the present invention include: acetylation,addition reactions, alkylation, dealkylation, hydrodealkylation,reductive alkylation, amination, aromatization, arylation, autothermalreforming, carbonylation, decarbonylation, reductive carbonylation,carboxylation, reductive carboxylation, reductive coupling,condensation, cracking, hydrocracking, cyclization,cyclooligomerization, dehalogenation, dimerization, epoxidation,esterification, exchange, Fischer-Tropsch, halogenation,hydrohalogenation, homologation, hydration, dehydration, hydrogenation,dehydrogenation, hydrocarboxylation, hydroformylation, hydrogeniolysis,hydrometallation, hydrosilation, hydrolysis, hydrotreating (HDS/HDN),isomerization, methylation, demethylation, metathesis, nitration,oxidation, partial oxidation, polymerization, reduction, reformation,reverse water gas shift, sulfonation, telomerization,transesterification, trimerization, and water gas shift.

Another advantage of the present invention is that good yields (or othermeasures of good performance) can be obtained with short contact times.In preferred methods, the contact time is less than 100 milliseconds(ms), more preferably less than 50 ms and still more preferably between1 and 25 ms for gas phase reactions. Liquid phase reactions would beexpected to be at least three orders of magnitude slower. Contact timemay be reduced by reducing the diffusion distance between the bulk flowand the porous catalyst while concurrently reducing channel length. Atthese contact times, in a preferred embodiment of hydrocarbon steamreforming, at least 70%, more preferably at least 90%, of theequilibrium conversion of the hydrocarbon entering the beginning of saidat least one reaction chamber is converted to hydrogen, carbon monoxideand/or carbon dioxide. Similar improvements can be obtained in otherprocesses.

Some process characteristics of some preferred inventive processesinclude the following:

-   -   1. Operate safely at a fuel:oxygen ratio near stoichiometric        (less than 100% excess air) for the use of combustion as the        exothermic reaction. This reduces the required air which        improves the overall system thermal efficiency and reduces the        required duty for the external air blower or compressor.    -   2. Operate steam reforming at short contact times or conversely        at high gas hourly space velocities. This is required to create        a compact device.    -   3. Operate with a high heat flux. This is required to operate at        short contact times.    -   4. Operate with a low pressure drop per unit length of reactor.        This enables a higher productivity per unit volume.    -   5. Optional: quench/inhibit gas phase reactions. As the channel        dimension nears the quench diameter or drops below, then the        contribution of the unwanted gas phase homogeneous combustion        reaction is reduced.

In its broader aspects, the invention relates to any pair (or more, thatis, different compositions can be run through different reactionchambers having different catalysts) of endothermic and exothermicreactions.

In steam reforming, gas hourly space velocity is preferably greater than10,000, more preferably greater than 50,000, and may range from about100,000 hr⁻¹ to over 10⁶ hr⁻¹ corresponding to a contact time on theorder of 100 to 1 milliseconds, respectively. Operating pressures formethane steam reforming may range from 1 atm to 50 atm. A range of 1 to30 atm is preferred. Steam to carbon ratios may range from 1 to 10; arange of 1 to 3 is preferred.

A variety of hydrocarbons can be reformed to produce hydrogen, includingmethane, ethane, propane, alkanes in general, alkenes, alcohols, ethers,ketones, and the like including blends and mixtures such as gasoline,diesel, kerosene, and others.

In addition, the present invention can be used to intensify otherendothermic reactions beyond steam reforming. As an example, the presentinvention could be used to intensify a dehydrogenation reaction bysupplying heat via an integrated combustion reaction.

EXAMPLES

Preferred catalysts for use in the apparatus described in the Exampleswere prepared according to the following procedures:

The catalyst in the reformer channels contained a catalyst of13.8%-Rh/6%-MgO/Al₂O₃ on a metal felt of FeCrAlY alloy obtained fromTechnetics, Deland, Fla. The reforming catalysts were prepared using awash-coating technique based on FeCrAlY felt with 0.01″ thickness and90% porosity. Before wash coating, metal felt was pretreated by a rapidheating to 900° C. in air for 2 hours. To enhance the adhesion betweenthe metal surface and the catalyst, a dense and pinhole-free interfaciallayer was first coated onto the oxidized FeCrAlY felt by metal organicchemical vapor deposition (MOCVD). This interfacial layer can be Al₂O₃,Al₂O₃+SiO₂, or TiO₂, etc. For example, when TiO₂ was coated, titaniumisopropoxide (Strem Chemical, Newburyport, Mass.) was vapor deposited ata temperature ranging from 250 to 900° C. at a pressure of 0.1 to 100torr. Titania coatings with excellent adhesion to the foam were obtainedat a deposition temperature of 600° C. and a reactor pressure of 3 torr.This layer not only increases the adhesion between metal felt and thecatalyst, it also protects the FeCrAlY from corrosion during the steamreforming reaction. 13.8 wt % Rh6 wt % MgO/Al₂O₃ powdered catalyst wasprepared by 1) calcining a high surface area gamma-alumina at 500° C.for 5 hours; 2) impregnating the gamma alumina with MgO using theincipient wetness method with an aqueous solution of magnesium nitrate;and obtaining an MgO modified gamma alumina support; 3) drying themodified support at 110° C. for 4 hours followed by 4) a secondcalcination at 900° C. for 2 hours; 5) impregnating the modified supportwith Rh₂O₃ with the incipient wetness method from a rhodium nitratesolution; 6) followed by a final drying at 110° C. for 4 hours and a 7)final calcinations at 500° C. for 3 hours to obtain a powder of thesupported catalyst. Catalyst coating slurry was prepared by mixingpowder catalyst aforementioned with de-ionized water in the ratio of1:6. The mixture was ball-milled for 24 hours to obtain coating slurrycontaining catalyst particles less than 1 micron. The heat-treated andCVD coated felt was wash-coated by dipping the felt into catalystslurry. The wash coating process may be repeated to obtain desiredweight gain. Between each coating, the felt coated with catalyst wasdried in an oven at 100° C. for 1 hour. The coating procedure isrepeated to achieve desired coating thickness or catalyst loading. Afterthe final coating step, the catalyst was dried overnight in an oven at100° C. and calcined by heating slowly in air at rate of 2° C./min to atemperature in the range of 300 to 500° C. The amount of catalyst coatedwas measured to be 0.1 gram catalyst per square inch (6.5 cm²) of felt.Prior to steam reforming testing, the engineered catalyst felt wassubjected to an activation treatment, preferably reduction at 300-400°C.

The integrated combustion catalyst can be a wash-coated catalyst that isapplied directly to the interior Inconel walls of the ICR device. TheInconel surface is first cleaned, ultrasonically if possible, in hexane,nitric acid (20%) and acetone (or propanol). Preferably, the cleaningsolutions are flowed over the Inconel surfaces. A native chromium oxidelayer is then formed on the Inconel surface by heating in air (flowing,if possible) at 3.5° C./min to 500° C., and held at 500° C. for 2 hours.The temperature is then increased at 3.5° C./min to 950° C., and held at950° C. for 2 hours. The Inconel is then allowed to cool to roomtemperature at a rate no faster than 5° C./min. The active palladiumcomponent is then applied to the chromia layer by submersing therequired deposition area in a 10-wt % solution of palladium nitrate.This is accomplished either by static submersion, or by pumping thefluid into a device to a required liquid level. The solution is thenallowed to remain in contact with the deposition surface for 2 minutes.The solution is then removed from contact with the Inconel surface, andthe amount of palladium remaining is calculated through a differencemeasurement. In the case of channel coating, nitrogen is flowed throughthe channel do ensure no plugging occurs. The catalyst is then dried at100° C. for one hour, under vacuum if possible. The catalyst is thencalcined by heating at 3.5° C./min to 850° C., held at 850° C. for 1hour. The catalyst is then allowed to cool to room temperature at a rateno greater than 5° C./min.

For some examples, a felt form of the combustion catalyst was preparedand then inserted into the combustion microchannel(s).

Engineered combustion catalyst was also prepared based on the FeCrAlYfelt from Technetics. Similar to the preparation of engineered steamreforming catalysts, the FeCrAlY felt substrate was first fired at 900Cin air for 2 h in a muffle furnace (ramping rate=20C/min). After thefiring process, the felt was cooled to room temperature. It was thendip-coated in a colloid Al₂O₃ solution (PQ corporation) containingmicron sized gamma Al₂O₃ particles. This step was conducted by immersingthe felt into the solution, then removing excess solution on the felt onan absorbent sheet, followed by drying under vacuum at 110C forovernight. The sample was heated to 500C for 3 h prior to the additionof Pd. The Pd was added by soaking the engineered substrate, nowcontaining an Al₂O₃ layer into a 20 wt % Pd(NO₃)₂ solution (Engelhard).Upon as removing excess Pd(NO₃)₂ solution, the sample was dried invacuum at 110 C for at least 4 h. Final calcination was conducted byheating at 2 C/min to 350 C, and holding isothermally at thattemperature for 3 hrs. The prepared Pd/Al₂O₃ engineered catalyst has anominal loading of 47 wt % Pd over Al₂O₃ and 0.126 g-cat/g of FeCrAlY.

Example 1

Integrated Combustor Reactor (ICR)

This integrated catalytic combustor reactor was composed of a singlemethane steam reformer channel that shares a wall with a singlecatalytic combustion channel. Heat was transferred through this commonwall from the hot (combustion) side to the cold (reforming reaction)side to drive the endothermic reaction. One design (version 1) of theICR had no header or footer space for the combustion stream, insteadintroducing the unmixed combustion gases directly onto the catalyst.This was done to insure that homogeneous combustion did not occurupstream of the catalyst (i.e. in the header). A second design of theICR (version 2) was fabricated which included a 0.25″ (6.4 mm) by 0.4″(10.2 mm) header and footer space on the combustion side, identical tothose on the reformer side (see FIG. 8). This was done to reduce thelikelihood of channeling on the combustor side and to more evenlydistribute the flow (and thus the heat of reaction) over the entirewidth of the reactor, thus prolonging catalyst life and increasingreformer conversion. The overall dimensions of each ICR unit were0.23″×0.5″×1.7″ and the dimensions of each catalyst were0.01″×0.5″×1.0″. For the purposes of heat flux calculations for thisexample only, 0.4″ catalyst width was included as the remaining catalystvolume was occluded from reactant flow. Both the reformer and combustorchannel catalysts operated with flow of the reactants through thecatalyst (and therefore excessively high pressure drop), having noflow-by gap once assembled. The experimental test conditions used areshown in Table 1. Combustion fuel and air were not preheated and SMRreactants were preheated to ˜600-650° C. Although outlet pressures arenot reported, they are close to ambient (˜1 bar absolute).

The integrated combustor reactor devices (version 1 and version 2) werefabricated from a 0.010 inch (0.25 mm) Inconel 625 heat transfer shimthat was perimeter welded between two 0.115 inch (2.9 mm) by 0.7 inch(17.8 mm) by 1.7 inch (43.2 mm) plates. One plate held the engineeredcombustion catalyst against the heat transfer shim, while the otherplate held the engineered SMR catalyst against the opposite side of theshim. Each plate admitted reactants (combustion or SMR) into a cavity(machined into the inside face) through a 0.125 inch (3.2 mm) O.D. (1.8mm I.D.) inconel tube at one end and discharged products through asimilar tube at the other. Except in the case of the version 1combustion plate, a 0.25 inch (6.4 mm) length of flow-path between thecatalyst and the inlet or outlet tube was included to provide for headerand footer regions to help distribute the flow throughout thecross-section. Each catalyst was held in intimate thermal contact withthe heat transfer shim by 0.050 inch (1.3 mm) by 1.0 inch (25.4 mm)rails machined on either side of the flow path. Fuel was introduced intothe device through a 0.063 inch (1.6 mm) tube (1.1 mm I.D.) nested in aconcentric fashion inside the air inlet tube, so that the air and fuelwere not allowed to mix until the end of the 1.6 mill tube (about 1-2 mmfrom the heat transfer shim).

The heat transfer area between the two catalysts was 0.4 inches (10.2mm) by 1.0 inches (25.4 mm). The reactor core volume (used to calculateaverage volumetric heat flux) was assumed to be the heat transfer areatimes the sum of die heat transfer web thickness (0.25 mm), SMR channelthickness (0.25 mm), and combustion channel thickness (0.25 mm).

After two thermal cycles and six hours of operation using the firstconfiguration, the flow configuration was changed from co-flow tocounter-flow in an attempt to further improve the heat transfer to thesteam reformer. Only hydrogen/air mixtures were used in the operation ofthe first design in order to eliminate the possibility of coke formationin the combustor.

Version 2 was tested continuously (in a co-flow configuration) for threehours combusting a hydrogen/air mixture and for four hours combusting amethane/air mixture. During operation of ICR version 2 the fuelequivalence ratio was not allowed to exceed 0.7, thus maintainingmaximum adiabatic flame temperatures for H₂/air mixtures of 1740° C. (asopposed to ˜2110° C. at stoichiometric) and 1565° C. for CH₄/airmixtures. TABLE 1 Conditions used in reduction to practice of ICR.Version 1 Version 2 Version 2 CR fuel/oxidant H₂/air H₂/air CH₄/air CRcontact time 1.75 ms 1.60-1.92 1.35-1.92 CR excess air 5% 43% 43% SRcontact time 14-30 ms 9.5-19.0 10.0-19.7 SR steam to carbon ratio2.8-3.4 3.1-3.2  3.0-3.1 flow configuration co/counter-flow co-flowco-flowCR refers to the combustion reaction,while SR refers to the steam reforming of methane reaction.

Results

Version 1

A hydrogen conversion of 99.9% was achieved in the integrated combustorreactor with 5% excess air (fuel equivalence ratio of 0.95), giving amaximum combustor exit temperature of 1050° C. or less. In contrast, anexternal hydrogen combustor would require 186% excess air (fuelequivalence ratio of 0.35) to keep the adiabatic combustion producttemperature to 1050° C. or below. A typical pressure drop through thecombustion channel was ˜40 psig for the 1.75 ms contact time. Pressuredrop in the microreactor with integrated combustor was much higher onthe combustion side due primarily to the much higher flow rates (muchshorter contact times) and secondarily to the manner in which the airwas fed (through a very narrow annulus). TABLE 1 Summary of best resultsfrom devices of example 1. Version 1 Version 1 Version 2 Version 2Version 2 Version 2 air inlet T (° C.) 25 25 25 25 25 25 air inletpressure (psig) 41 41 59 59 78 78 air inlet pressure (Pa/10⁵) 3.8 3.85.1 5.1 6.4 6.4 fuel inlet T (° C.) 25 25 25 25 25 25 CR^(a) H2 flow(SLPM) 0.326 0.326 0.304 0.304 0 0 CR^(a) CH4 flow (SLPM) 0 0 0 0 0.0990.099 CR air flow (SLPM) 0.78 0.78 0.975 0.975 1.385 1.385 excess air  0%   0%   35%   35%   47%   47% CR max. meas. temperature (° C.) 10551046 936 902 864 839 CR contact time (ms) 3.6 3.6 3.1 3.1 2.7 2.7 CRGHSV (per hour) 1000000 1000000 1161290 1161290 1333333 1333333 CR H₂conversion NM 99.9% 99.9% 100.0%  99.6% NM CR pressure drop (psi) <41<41 <59 <59 <78 <78 CR pressure drop (Pa/10⁵) <3.8 <3.8 <5.1 <5.1 <6.4<6.4 SR preheat T (° C.) 645 645 ˜580 ˜580 ˜580 ˜580 SR inlet pressure(psig) 10.5 24.5 44.5 64 48 72 SR inlet pressure (Pa/10⁵) 1.7 2.7 4.15.4 4.3 6.0 SR^(a) CH₄ flow (SLPM) 0.0167 0.0485 0.0253 0.0498 0.02530.0498 SR H₂O flow (SLPM) 0.05 0.152 0.0785 0.158 0.0748 0.148 steam:C(mol:mol) 3 3.1 3.1 3.2 3 3 SR contact time (ms) 60 20 38 19 39 20 SRGHSV (per hour) 60000 180000 94737 189474 92308 180000 SR bodytemperature (° C.) 790 773 812 785 840 813 SR CH₄ conversion 79.7% 54.5%97.7% 93.6% 98.0% 94.0% SR selectivity to CO 79.5% 66.6% 61.0% 60.2%66.3% 65.4% SR pressure drop (psi) <10.5 <24.5 <44.5 <64 <48 <72 SRpressure drop (Pa/10⁵) <1.7 <2.7 <4.1 <5.4 <4.3 <6.0 avg. heat flux(W/cm²) 0.8 1.6 1.5 2.7 1.5 2.8 avg. volumetric heat flux(W/cm{circumflex over ( )}3) 10.8 20.7 19.1 35.7 19.5 36.6^(a)CR refers to combustion process side; SR refers to reforming processside.

The ICR steam reformer (for a steam to carbon ratio of ˜3) was able toextract heat from the integrated combustor in the co-flow configuration,maintain steam reforming temperatures of 780-800° C. and methaneconversions as high as 75%, with reformer contact times of 20-60 ms (seeFIGS. 9-10, contact time calculated based on entire channel volume,including catalyst and the total gas flowrate at standard temperatureand pressure).

It was found that, in this device, for combustor whole channel contacttimes of less than 3.5 ms (5% excess air), the hydrogen combustion zoneappeared to extend past the exit of the catalytic combustion area. Adecreasing combustor exit gas temperature and an increasing gastemperature immediately downstream of the exit with increasing combustorflow rate evidenced this. High combustor fuel flow rates (i.e. <4 mscontact times) were required to achieve reformer temperatures of 750° C.and higher. This is likely due to channeling of the combustion streamdown the central 30% of the combustion catalyst. If channeling andcatalyst deactivation could be eliminated, it is thought thathydrogen/air contact times of <1.2 ms might be achieved. Althoughcombustor exit temperatures were far below adiabatic flame temperatures,examination of the combustion catalyst revealed evidence of noble metalsintering or vaporization and re-deposition. This suggests that anunacceptably high-temperature region existed in the combustor (>1200°C.).

Version 2 Combustor (CR)

The inclusion of a header in the combustor of ICR version 2 greatlyimproved the efficiency of operation of both the combustor and thereformer. Results of combustor tests for ICR version 2 are shown inTable 2. The highest pressure drop across the combustor side of the ICRversion 2 was 5.4 bar (78 psid), for CH₄/air at 2.8 ms whole channelcontact time.

During operation of the version 2 device, two combustor operationregimes were observed. In the first regime, the maximum temperature seenon the steam reforming side was in the header, directly opposite thecombustor inlet (typically >100° C. above the maximum body temperature).This is strong evidence that, in this regime, a homogeneous combustionzone exists in the header of the combustor. In the second combustionregime, the maximum temperature in the reformer is the body temperature(at the thermocouple well in the center of the side plate, sunk 1.78 mmdeep into tile plate wall), rather than in the header. In addition, inthe second regime the temperature at the combustor inlet drops below theauto-ignition temperature for the fuel/air mixture. This is evidencethat regime 2 does not involve combustion in the header. In other words,in regime 2 a flame will not attach in a stable fashion to the fuelinlet tube in the header. Table 2 describes the conditions for whicheach regime was observed during operation of version 2. The transitionbetween regime 1/regime 2 (as identified by a sudden discontinuouschange in the reformer header temperature) was observed twice duringtesting, each time ˜3-4 minutes after a change was made in the processconditions.

In this device at the conditions shown in Table 2, hydrogen combustionoperates only in regime 1, probably due to the unusually high flamespeed of hydrogen/air mixtures (an order of magnitude higher thanmethane/air mixtures). On the other hand, methane combustion wasobserved in both regimes, with the transition occurring somewherebetween ˜600-1000 sccm total reactant flow. Both combustionfuels/regimes gave similar results (in terms of steam reformerperformance) at identical total flows and fuel equivalence ratios,although reformer temperatures were much more uniform (along the flowpath) for regime 2. No decrease in combustor fuel conversion wasobserved over the course of seven hours of operation. TABLE 2 Results ofcombustor analysis during operation of ICR unit at a fuel equivalenceratio of 0.7 (43% excess air). fuel contact body CO₂ flow air flow timere- T^(a) conversion select. fuel (sccm) (sccm) (ms) gime (° C.) (%) (%)H₂ 228 792 1.9 1 785 100.0  — H₂ 280 951 1.6 1 810 97.9-100.0 — CH₄ 70956 1.9 2 775 99.1 100.0 CH₄ 84 1148 1.6 2 810 99.4 100.0 CH₄ 99 13561.4 2 860 99.6 100.0 CH₄ 43 570 3.2 1 regime transition during shut down^(a)Combustor body temperatures are for center thermocouple well andcorrespond to baseline reformer case (100 sccm SR flow). In every casethis body temperature was the highest temperature observed on thecombustor side (although higher internal temperatures were clearlypresent) and the footer temperature was the lowest observed.Version 2 Steam Reformer (SR)

Because of the ability of the version 2 combustor to perform moreefficiently and more uniformly, the steam reformer was able to achievemuch higher conversions than the version 1 reformer. The results of theversion 2 ICR are shown in FIGS. 11-12. As expected, conversionincreases with increasing temperature (see FIG. 11), but the temperatureinside the reformer is difficult to ascertain in such a small device, asthe header, footer, and body temperatures are likely much lower than theinside gas temperature. Conversion of SR methane decreased withdecreasing contact time between 40 and 20 ms (see FIG. 12), although thechange is exaggerated since the temperature also dropped with increasingthroughput (by at least 25° C.). Selectivity to CO increasessignificantly with increasing temperature (FIG. 11), but not withincreasing contact time (FIG. 12). The highest SR conversion was at thehottest condition (65.4 W_(thermal) methane combusted in 43% excessair), where 98.6% methane conversion and 66% selectivity to CO wereachieved at 48 psig feed pressure, 3:1 steam:carbon, and 38.0 ms (SR)contact time. After 1.5 hours of operation at the hottest combustortemperature, the conversion dropped to 98.0%, evidence of slightdeactivation. The lowest SR methane conversion was 93.6% for a SRcontact time of 19.0 ms. The steam reformer was on line for a total of6.5 hours. The carbon balance closed to within ±4% for all but onecondition.

Conclusions

The following are significant findings of Example 1:

-   -   1. Integrated hydrogen/air combustion (flow-by configuration)        looks very promising for well mixed, uniformly dispersed streams        low excess air, and whole channel contact times as low as 3.2 ms        if a header is included in the design. 100% H₂ conversion was        achieved under these combustor (CR) conditions. Integrated        methane/air combustion (flow-by configuration) looks extremely        promising for well mixed, uniformly dispersed CR streams with        low excess air, and whole channel contact times as low as 2.7 ms        if a header is included in the design. An incredible 99.6% CH₄        conversion with 100% selectivity to CO₂ was achieved under these        CR conditions, with no air preheat. Such performance could        require at least a contact time five times longer (and about        400° C. of air preheat) for methane combustion in an external        catalytic combustor.    -   2. An integrated catalytic combustion channel can provide the        heat necessary to sustain an endothermic catalytic steam        reforming reaction in an adjacent channel of the same size even        under the “worst case” conditions of heat loss existing in the        ICR version 2 device. SR methane conversions as high as 98.6%        with 66% selectivity to CO were observed.    -   3. Insufficient dispersion of combustion reactants across the        combustion catalyst cross section (i.e. channelling) results in        greatly reduced conversions in both the combustor and reformer.    -   4. The design of these ICR devices (versions 1 & 2) caused an        unexpectedly large pressure drop in both the reformer (˜3.3 bar        or 48 psid) and combustor (˜5.4 bar or 78 psid) sections when        operated at temperature (˜850-900° C.) with a 40 ms SR contact        time.

Example 2

This example describes the design, fabrication, and test results from ahigh efficiency, high-throughput small microchannel reactor in whichheat producing (exothermic) and heat consuming (endothermic) reactionchannels are immediately adjacent (integrated). Combustion of hydrogenin air was used as the exothermic reaction, while steam reforming ofmethane with a steam to carbon ratio of 3:1 was used as the endothermicreaction. A new ICR design (flow-by) was used which allowed for muchhigher throughput with minimal (i.e. <11 psi) pressure drop by allowingeach reactant stream to flow in a narrow (0.125 mm) gap adjacent to theporous engineered catalyst. The new design included a central combustionzone (of two microchannels) flanked by a reformer channel on eitherside.

The ICR (see FIG. 13) used a shortened combustion catalyst bed (0.2″,0.5 cm) to allow a combination of catalytic and homogenous hydrogencombustion. Hydrogen was only distributed across the channel width, andnot along the length of the catalyst bed. The entire hydrogen fuelstream was fed into the combustion channel through eight 0.009″ (0.02cm) holes 0.030″ (0.08 cm) upstream of to the catalyst bed.

The dimensions of the steam reformer (SR) catalysts were 0.01″×0.5″×1.0″(0.25 mm×1.27 cm×2.54 cm) and the combustor (CR) catalysts were0.01″×0.5″×0.2″ (0.25 mm×1.27 cm×0.51 cm), although only 1.02 cm of thewidth was exposed to reactant flow, 0.12 cm on each edge being used tohold the catalyst in place on either side. For the purposes of heat fluxcalculations for this example, only 0.4 inches catalyst width wasincluded as the remaining catalyst volume was occluded from reactantflow. A thickness of 0.012″ (0.30 mm) was allowed for each catalyst,with flow-by channel thicknesses of 0.005″ (0.13 mm) and 0.017″ (0.43mm) in the SR and CR channels, respectively. The ICR was operated in aco-flow configuration. The experimental setup used for testing was thesame as that used in Example 1. The ICR was tested with both astoichiometric amount of hydrogen in air and with 40% excess air in theCR channels while running methane steam reforming reaction in the SRchannels.

Body temperatures were measured inside 1/16″ (0.16 cm) deep thermowellsat 3 positions along the length of the ICR unit on the back side of eachreformer channel at positions corresponding to the top, middle, andbottom of the steam reformer catalyst bed. The combustor was tested fora range of contact times and air equivalence ratios using hydrogen inair supplied by Brooks mass flow controllers (MFC). The reformer wastested over a wide range of reformer contact times at an averagetemperature of about 750-800° C. Methane was supplied by a Brooks MFCand water was supplied by HPLC or syringe pump. Exit compositions weredetermined for both the combustor and reformer portions using an MTI GC.All tests were performed at ambient pressure using a 3:1 molar steam tocarbon ratio.

A summary of the results of tests performed is shown in Table 3. Thetests show a large drop (36-80° C.) in SR body temperatures along thelength of the reactor, suggesting that the majority of the combustionoccurs at the entrance of the combustion chamber (in the catalyst zone).It is noteworthy that the ideal region to supply heat to the reformingreaction is at the catalyst entrance, where the SMR reactantconcentration gradient (and therefore the rate of reaction) is highest.The co-flow configuration of the ICR and SMR channels coupled with thefast kinetics of hydrogen combustion delivers the majority of the heatright at the catalyst bed entrance, thus maximizing the rate of SMRreaction. TABLE 3 Summary of best results of tests using ICR device ofexample 2. 1 2 3 4 CR H2 flow (sccm) 316 632 767 857 CR air flow (sccm)1028 2092 2682 2682 Air inlet temperature (° C.) 25 25 25 25 Air inletpressure (psig) 2.0 4.9 6.7 6.7 excess air   37%   39%   47%   31% CRcontact time (ms)^(a) 3.4 1.7 1.3 1.3 CR flow time (ms)^(b) 17.0 8.4 6.66.4 CR H₂ conversion 92.8% — 92.3% — CR pressure drop (bar) 0.14 0.340.46 0.46 SR inlet temperature (° C.) 622 544 571 511 SR inlet pressure(psig) 1.4 6.6 10 11 SR CH₄ flow (sccm) 55 231 462 462 SR H₂O flow(sccm) 149 673 1346 1346 steam:C (mol:mol) 2.7 2.9 2.9 2.9 SR contacttime (ms)^(a) 65 14.8 7.4 7.4 T₁/T₄ (° C.)^(c) 770/820 768/859 705/813715/825 T₂/T₅ (° C.)^(c) 736/777 756/812 686/754 696/767 T₃/T₆ (°C.)^(c) 731/757 754/793 683/733 694/745 SR CH₄ conversion 98.3% 99.2%89.2% 93.4% SR selectivity to CO 65.0% 65.9% 61.0% 61.2% avg. heat flux(W/cm²) 1.7 7.1 12.6 13.2 avg. volumetric flux (W/cm³) 6.0 25.0 44.346.4 SR pressure drop (bar) 0.10 0.45 0.69 0.76 SR carbon balance  2.6% 5.2% −0.7% −8.8%^(a)Contact time based only on volume of channel and catalyst in regioncontaining catalyst, 0° C., 1atm.^(b)Flow time based on entire combustion chamber volume, includingdownstream of catalyst.^(c)Thermocouples are located down the length of the reactor, with thehighest temperatures near the inlet of the co-flow reforming stream andcombustion stream

The device showed differences of over 50° C. in SR body temperatures atidentical locations on each reformer channel, suggesting that the fuelor air is not uniformly distributed between the two combustorhalf-channels. The ICR device transferred tip to 54.8% of the heatproduced by the combustion into the endothermic reaction (it isestimated that about 40% of the energy produced was spent to heat thecombustion products to 740° C.). Thus when operated at peak capacity,heat losses in the device were as low as 10% of the energy available at740° C., demonstrating an incredibly high efficiency for such a smalldevice. The calculated maximum average heat flux was 13.2 W/cm². Thisflux is higher than any reported in the literature for driving anendothermic reaction. It follows, then, that the flow-by ICR concept canbe used to make high-throughput reactors which are much more compact andbetter performing than devices which use monolith or flow-throughcatalyst configurations.

Results from tests on the second type of integrated combustion reactor(ICR) show much higher efficiency than is possible with an externalcombustor (as high as 56.7% of combustion fuel energy were transferredto the steam reforming reaction, of the maximum possible of 60.4%deliverable at 740° C. with 92% combustion fuel conversion and 47%excess air). The heat flux was 13 W/cm², which is higher than anyreported in the literature. For the conditions at which this heat fluxoccurred, the pressure drops measured across the combustor and reformerwere 0.46 bar (6.7 psid) and 0.76 bar (11.0 psid), respectively. Theseresults confirm that the flow-by catalyst configuration enablescatalytic combustion heat transfer rates and efficiencies far abovethose in existing ICR-like devices. In contrast, the flow through(Example 1) device incurred pressure drops across the combustor andreformer of 5.3 bar (77 psid) and 5.0 bar (72 psid) at a heat flux of3.0 W/cm².

Results typical of the reformer performance are shown in FIG. 14. Valuesof equilibrium selectivity to CO calculated for measured reactortemperatures (assumed to be equal to the average of all six bodytemperatures) were in good agreement with measured selectivity values.From the data in FIG. 14 it is clear that the ICR is able to achieveequilibrium conversion for reformer contact times of 15 ms, and nearlyequilibrium conversion for a contact time of 7.5 ms.

FIGS. 15-16 support the hypothesis that either the fuel or the air wasnot evenly distributed between the two half-channels. The data in FIG.15 show that, for the range of contact times tested, no decrease inconversion is seen as contact time is shortened, suggesting that thelimitation on conversion is not due to catalytic activity.

A strong trend in hydrogen conversion is seen with decreasing excessair, suggesting that below about 150% excess air, one of the combustionhalf channels may be running fuel rich, limiting the maximum achievableoverall conversion.

The ICR was operated for a total of about 21 hours over 6 days, with nosignificant deactivation observed for either the combustor or reformercatalysts. During testing, the measured body temperatures never exceeded860° C. and were typically 750-800° C.

Conclusions

-   1. An integrated combustor reactor (ICR) device was designed,    fabricated and shown to be able to achieve a rate of heat transfer    of 13.2 W/cm² with pressure drops of less than 0.76 bar (11 psid)    and 0.46 bar (7 psid) in the reformer and combustor channels,    respectively.-   2. Equilibrium conversion and selectivity were achieved in this ICR    device for whole channel reformer contact times as low as 15 ms,    even though the combustion catalyst bed extended only ⅕ the length    of the reformer catalyst bed.-   3. After 21 hours of operation at 750-800° C., the single channel    ICR device showed no significant signs of catalyst deactivation,    even though the combustor was normally run with only 41% excess air    (with some excursions to <3% excess air).    Because of the unusually high capacity of this “single channel” ICR    device, heat losses as low as 10% of the total energy available at    740° C. were achievable.

Example 3

This example describes the design and testing of a multichannelinterleaved microchannel reactor in which heat producing (exothermic)and heat consuming (endothermic) reaction channels are interleaved(integrated) with each other. Combustion of hydrogen in air was used asthe exothermic reaction, while steam reforming of methane with a steamto carbon ratio of 2:1 was used as the endothermic reaction. Thepressure for both reactions was near atmospheric and operated at apressure required to overcome the system pressure drop. For the SMRside, the typical operating pressure was 139 kPa (5.5 psig) and for thecombustion side the typical operating pressure was also 137 kPa (5.2psig). A flow-by design (flow-by) was used which allowed for much higherthroughput with minimal (<0.2 bar, or 3 psid) pressure drop by allowingthe reactant stream in each channel (whether exothermic or endothermic)to flow in a narrow (about 0.2 mm) gap between two layers of porousengineered catalysts which are in intimate thermal contact with thesolid channel wall or heat transfer surface. The reactants substantiallydiffuse from the gap to the porous engineered catalyst. The reactantsthen continue to diffuse and react within the porous engineered catalystthat is adjacent to the heat transfer surface. The design included ninecombustion channels interleaved between ten reformer channels in aco-flow arrangement. The integrated combustion reactor (ICR) testresults demonstrate that the ICR concept can be scaled up for use in amultichannel device.

An ICR device body was fabricated from Inconel 625 by using wire EDM toform slots in a block of solid metal. The device was made from a 5.33 cm(2.1″) long by 5.23 cm (2.06″) tall by 2.54 cm (1.0″) wide block. Ninecombustion channels were machined 5.3 cm through the block's length bywire EDM, with each channel having an overall width of 1.78 cm (0.7″)and 0.081 cm (0.032″) height. Each combustion channel was centered inthe 2.5 cm width of the block. Interleaved between the nine combustionchannels were eight SMR channels, and bracketing these interleavedchannels were two smaller SMR channels, ensuring a heat sink on bothsides of each combustion channel. All the SMR channels in the block were4.57 cm (1.8″) in length, with the channel starting 0.25 cm (0.1″) fromthe side combustion reactants enter and 0.51 cm (0.2″) from the exit ofthe combustion channels. The eight SMR channels had heights of 0.081 cm(0.032″), while the two smaller bracketing SMR channels had heights of0.043 cm (0.017″). The separation between all SMR and combustionchannels was maintained by heat transfer webs of 0.15 cm (0.060″)thickness, and the interleaved channels were centered within the 5.23 cmof the device's height. The general channel orientation and theinterleaved nature of the channels are both illustrated in FIG. 17.

Fluid access to the SMR channels is through the sides where the channelsare machined through in a direction perpendicular to the combustionchannels. To cause the SMR flow in each channel travel concurrently withthe combustion channels, two 4.01 cm (1.6″) by 0.38 cm (0.15″) sidepieces were made for each SMR channel with heights matching the SMRchannel heights (press-fit). For each SMR channel, one of these sidepieces was placed 0.38 cm into the channel and positioned such that a0.51 cm opening was left upstream of the catalyst for reactant gas topass through from the header, while another was placed on the oppositeside, leaving a similar opening downstream of the catalyst to allowgases to leave the footer. These openings start 0.25 cm from thecombustion inlet side of the block. The flow-by SMR engineered catalystsare then inserted into the SMR channels, of length 3.57 cm (1.4″), 1.78cm (0.7″) width and 0.03 cm (0.012″) thickness. Two SMR engineeredcatalysts go into each of the eight center channels up against both heattransfer surfaces, resulting in a gap of is 0.02 cm (0.008″). A singleSMR engineered catalyst goes into the two bracketing SMR channels,places against the heat transfer surface it shares with the combustionchannel, leading to a 0.012 cm (0.005″) gap. The catalysts are centeredwithin the 4.57 cm length and 2.54 cm width of each SMR channel. Tomaintain the flow-by gap mentioned above and keep the engineeredcatalysts against the channel wall, two to 3.57 cm (1.4″) length, 0.13cm (0.05″) width Inconel 625 strips were placed between the twoengineered catalysts, with the thickness of the two strips equal to thegap height design. These two strips are aligned with their lengthscorresponding to the engineered catalyst lengths, and are placed at theedges of channels the two engineered catalysts form. The final stepinvolves placing the other side piece in 0.38 cm into the width of eachSMR channel on the opposite of the first piece. The placement creates a0.51 cm length opening for SMR products to leave the channel, and thisopening is 0.51 cm length from the edge that combustion products leavethe device. The overall effect of the addition of the two side are tobuild a z-shaped manifold for flow through the channel created by theSMR catalyst gap. A drawing that illustrates the shape of the flow pathfor the SMR channels is found in illustration 18.

The SMR manifold that distributes the SMR streams from a single inletpipe to the ten overall channels is made from a single block of material5.08 cm (2.0″) long, 4.71 cm (1.858″) tall and 0.95 cm wide (0.375″). Inthis block individual circular cylinder channels lead from the openingsmentioned in the preceding paragraph to a half tube welded over theopening. The SMR manifold that takes the outlets from each stream into asingle outlet tube has a similar design, with its connection cylindersmatching to the outlet channels.

The nine combustion channels each had two combustion engineeredcatalysts of the same dimensions as the SMR catalysts, and used the samedimensions for the strips used for separation in the SMR channels. Thecombustion engineered catalysts are pushed into the length or eachchannel 0.76 cm (0.3″) from the side the combustion reactants enter, andby doing so ensure the SMR and combustion engineered catalysts surfaceareas are aligned for optimal heat transfer.

The combustion gas manifold that brings in the fuel (hydrogen) and airmust distribute air and fuel equally to the nine channels while keepingthe two combustible streams separate until meeting at the combustionchannel entrance. This is accomplished by taking a 1.27 cm (0.5″)×5.23cm (2.058″)×2.54 (1″) cm wide block of Inconel 625 and using a wire EDMto machine 1.27 cm×1.78 cm×0.03 cm through-channels, one for the air andfuel for each channel. A single air channel lines up with the entranceof each combustion channel, while a single fuel channel is a height0.193 cm (0.076″) beneath it. The two channels connect at the combustionchannel entrance by use of a 0.038 cm (0.015″) length channel machinedinto the plane that interfaces the combustion channel. This allows thefuel to make a 90° turn and travel in the machined-in pocket at theinterface with the main body to meet the incoming air steam, wheremixing occurs, as noted in FIG. 18.e The fuel and air channels thus madefrom wire EDM are closed with welds on the surface opposite theinterface with the reactor body. The manifold for delivering air from asingle inlet tube into the wire EDM created channels is created bymachining in 0.25 cm (0.100″) diameter tubes that extend in from theside of the block. These tubes extend in 2.16 cm of width from the side,creating an individual manifold for air across the combustion channel.These individual channel manifolds are connected with a single tube tothe inlet. The manifold for the fuel has a similar design. The overallmanifold is shown in FIGS. 17 and 18.

The manifold for the combustion product to a single outlet tube was madefrom an Inconel 625 block of 0.95 cm long (0.375″), 4.71 cm (1.858″)tall and 2.29 cm wide (0.9″). The manifold is a large open space of0.200″ length for flow to collect into, with a 0.46 cm (0.18 ″) diameteroutlet on the bottom side of the face opposite tile interface with thereactor.

Gaseous reactants were metered from pressurized cylinders via Brooksmass flow controllers, while water was delivered via high pressureliquid (HPLC) pump). An external combustor and heat exchange network wasused to continuously vaporize the feed water, preheat the steamreforming reactants to 500-600° C., and to preheat the combustion air to500° C. Device body temperatures were measured in thermowells (drilledinto selected webs between channels) using 0.032″ type K thermocouplesand recorded using an Iotech data logging system. Inlet and outlet gaspressures were measured for the combustion and reforming streams viacalibrated pressure transducers. Gas temperatures were also measured viathermocouple at the combustion gas inlet and outlet and at the reformergas inlet and outlet. The product gases were cooled in water-cooledcondensers and the dry exit flow measured in a dry test meter. Dryproduct gas compositions (H₂, CO, CO₂, CH₄, O₂, and N₂) were measuredusing an online gas chromatograph. All device ICR tests were performedusing a 2:1 molar steam to carbon ratio.

Catalysts were reduced prior to the first day of testing at 125° C. for1.5 hours. Startup of tile device ICR was accomplished in several steps.First an inert flow (nitrogen, 1-5 SLPM) was initiated in both thecombustion and reforming channels. Next the external combustor wasignited to provide water vaporization and preheat in externalmicrochannel based heat exchangers. Then the nitrogen flow to thecombustor was shut off and the air flow initiated. Hydrogen was thenintroduced to the combustion channels as about 3% of the total flow,causing catalytic ignition of the combustion reactants. Next thehydrogen content was slowly ramped up to bring the reactor temperatureto about 700° C. Then a small hydrogen flow was introduced into thereforming channels to prevent catalyst oxidation by steam, water was fedthrough the vaporizer, thus providing steam to the reforming channels.Finally, the reformer nitrogen was turned off and methane was introducedinto the reformer and the reformer hydrogen was turned off. Once thereforming reaction began, additional hydrogen was fed to the combustionchannels to maintain the reactor temperature between 750 and 850° C.Changes in process flows made during steady-state operation wereperformed in a balanced fashion (changing both the reforming andcombustion flows at the same time) to prevent large transient deviationsof the device ICR temperature from the desired operating temperature.

A summary of the results of tests performed using the device ICR at nearambient pressure is shown in Table 4. Note, in Table 4, that since thecombustion air was preheated to 500° C. and a low percentage of excessair (48-75%) was used, adiabatic flame temperatures inside thecombustion channels would theoretically exceed 2100° C. However, maximummeasured temperatures inside the device ICR were well below 950° C.,demonstrating the effectiveness of the ICR concept in transferring heataway from the combustion channels and into the endothermic reaction.Also note in Table 4 that high heat fluxes (nearly 3 W/cm²) and highconversions can be achieved even at very short contact times (as low as33 ms). Contact time is defined here as the total channel volume (withinwhich the catalyst and flow-by area are contained) divided by the totalvolumetric inlet flow rate (converted to standard conditions, 0° C. and1 atm). Even for the very fast contact times of the data shown in Table4, pressure drop is minimal (about 0.14-0.21 bar), demonstrating thedistinct advantage of the flow-by arrangement in the ICR. TABLE 4Summary of best results for high-throughput tests using ICR device ofEx. 3. Experiment 1 Experiment 2 air inlet T (° C.) 493 494 air inletpressure (psig) 5.2 4.16 air inlet pressure (Pa/10⁵) 1.4 1.3 fuel inlettemperature (° C.) 25 25 CR^(a) H2 flow (SLPM) 2.4 3.4 CR^(a) CH4 flow(SLPM) 0 0 CR air flow (SLPM) 10 12 excess air   75%   48% CR max. meas.temperature (° C.) 916 903 CR contact time (ms) 22.4 18 CR GHSV (perhour) 160714 200000 CR H₂ conversion 99.4% NM CR pressure drop (psi) 2.33.2 CR pressure drop (Pa/10⁵) 1.2 1.2 SR preheat T (° C.) 515 568 SRinlet pressure (psig) 3.0 5.5 SR inlet pressure (Pa/10⁵) 1.2 1.4 SR^(a)CH₄ flow (SLPM) 1.54 2.79 SR H₂O flow (SLPM) 3.11 5.61 steam:C (mol:mol)2 2 SR contact time (ms) 60 33 CR GHSV (per hour) 60000 109091 SR bodytemperature (° C.) 780 770 SR CH₄ conversion 82.3% 70.7% SR selectivityto CO 76.3% 77.6% SR pressure drop (psi) 1.8 3.3 SR pressure drop(Pa/10⁵) 1.1 1.2 avg. heat flux (W/cm²) 1.8 2.8 avg. volumetric heatflux (W/cm{circumflex over ( )}3) 8.9 13.5^(a)CR refers to combustion process side; SR refers to reforming processside.

This ICR device was designed to show proof of principle of the ICRconcept in a multiple channel device and is not an optimized prototype.Optimization of inlet and outlet manifolds, as well as channel geometry,is expected to allow increased capacity and/or decrease the pressuredrop penalty. Measurements of methane conversion in the device ICR(reformer channels) were conducted for a variety of combustion channelair equivalence ratios. Air equivalence ratios are defined as the ratioof available combustion oxygen to that required for stoichiometriccombustion. It was observed that methane conversion gradually decreasedwith increasing air equivalence ratio from about 83% conversion at a1.75 air equivalence ratio to about 63% conversion at a 7.0 airequivalence ratio. Although some coking of the reforming catalyst wasobserved when reforming reactant preheat was allowed to drop below 500°C., activity was completely restored after burning out the coke with airand re-reducing the catalyst. The ICR device was operated for over 10hours including 3 thermal cycles.

In summary, a multi-channel integrated combustor reactor device wasdesigned, fabricated and shown to be able to achieve a rate of heattransfer of 2.8 W/cm² and a volumetric heat flux of 13 W/cm3 withpressure drops of less than 0.23 bar and 0.19 bar in tie reformer andcombustor channels, respectively. High conversion (70-82%) was achievedin this ICR device for whole channel reformer contact times of 33-60 ms.After >10 hours of operation at 750-850° C., the device ICR deviceshowed no significant signs of catalyst deactivation, even though thecombustor was run with 75% excess air or less for more than 1.5 hours.

Example 4

An integrated-combustion microchannel reactor capable of providingexceptionally high heat fluxes (up to 29 W/cm² and 118 W/cm³) to ahighly endothermic reaction with minimal pressure drop was designed,fabricated, and demonstrated. This report describes the design andtesting of this multichannel interleaved microchannel reactor in whichheat producing (exothermic) and heat consuming (endothermic) reactionchannels were interleaved (integrated) with each other. Combustion ofhydrogen in air was used as the exothermic reaction, while steamreforming of methane (SMR) with a steam to carbon ratio of 2:1 was usedas the endothermic reaction. A flow-by design was used which allowed formuch higher throughput with minimal (<0.34 bar) pressure drop byallowing the reactant stream in each channel (whether exothermic orendothermic) to flow in a narrow (about 0.3 mm) gap between two layersof porous engineered catalyst which are in intimate thermal contact withthe solid channel wall or heat transfer surface. The design includedfive combustion channels interleaved between six reformer channels in aco-flow arrangement. The test results demonstrate that microchannelintegrated combustion devices can be used to deliver much highervolumetric heat fluxes (i.e. 118 W/cm³) than is possible in conventionalreformers and these heat fluxes can be achieved with minimal pressuredrop (0.28-0.34 bar). High conversions were also possible at very highgas hourly space velocities (e.g., 10⁶ hr⁻¹) were achieved.

Gas hourly space velocity is defined as the number of SMR reactor coregas volumes (defined at 0° C. and 1 atm) that would pass through the SMRcatalyst containing channels each hour. The space velocity can becalculated by dividing 3.6×10⁶ by the contact time in milliseconds, bothbeing defined at STP 0° C. and 1 atm. The reactor volume for use in allvolumetric heat flux calculations includes the entire core reactorvolume that is inclusive of all reformer channels including catalysts,all metal webs between reformer and combustion channels, and allcombustion channels used to supply heat. The external packaging is notincluded as this volume does not contribute to the transfer of heatbetween the two fluid streams.

Experimental Method

A high heat flux multi-channel reactor device with integrated combustionwas designed, fabricated, and tested. The device body was fabricatedfrom Inconel 625 using wire EDM to form slots in a 1 inch by 1.37 inchby 1.7 inch (2.5 cm×3.48 cm×4.3 cm) solid block of metal. External coverplates containing the inlet and outlet gas manifolds (i.e. headers andfooters) were independently machined and welded onto the body. Eachheader was designed to distribute the flow evenly from channel tochannel and within each channel. A drawing of the exterior of thisdevice is shown in FIG. 19 a including air inlet 102, hydrogen inlet104, reformer feed inlet 106, reformer product outlet 108, and combustoroutlet 110. For testing purposes, temperature monitoring wasaccomplished via thermocouples 112. All the small tubes 112 in FIG. 19are thermocouple tubes. FIGS. 19 b and 19 c show exploded views of theintegrated reformer. Air is fed into header 114 and distributed at equalpressure into slots 116 at equal pressure through each slot. Similarly,hydrogen is fed through inlet 104, and at equal pressure into each tube,into slotted tubes 118. The tubes 118 extend from the hydrogen inletinto manifold body 120. Schematic views of a gas mixer 125 is shown inFIGS. 19 d and 19 e (which are not drawn to scale). Projections 127define a gap 129 for air flow. Hydrogen exiting from slots 131 passesthrough channel 133 and mixes with air in combustion chamber 122 insidereactor body 124. As can be seen, this design mixes air and fuel insidethe combustion chamber and very near combustion catalyst 137; thusavoiding heat loss from combustion outside the combustion chamber. Steamreformer feed is fed into header 130 and passes, at equal pressure, intothe front of the steam reformer chambers. An illustration of thereformate flow through an individual chamber 132 is shown in FIG. 20.Plates 134 have supports 136 that press the reformer catalyst againstthe top and bottom of the reformer chamber.

The main body of the device of example 4 was identical to the main bodyof the device of example 3 except that there were four fewer SMR andfour fewer combustion channels, and the catalyst length in the flowdirection was only 1.0 inch rather than 1.4 inches. The dimensions ofthe SMR slots were 0.035 inches (0.89 mm) thick by 0.7 inches (17.8 mm)wide by 1.0 inch (25.4 mm) long. There were 6 SMR slots. The dimensionsof the combustion slots were 0.038 inches (0.97 mm) thick by 0.7 inches(17.8 mm) wide by 1.0 inch (25.4 mm) long. There were 5 combustionchannels. The channels were interspersed such that the outermost channelwas a reforming half channel adjacent to a combustion channel, thenadjacent to a full reforming channel and so on, in the same fashion asthe device of example 3. The reforming half channel was 0.018 inch (0.46mm) thick by 0.7 inches (17.8 mm) wide by 1.0 inch (25.4 mm) long. Thehalf channel only contained one catalyst insert that was placed next tothe wall shared with the combustion channel. The full reforming channelscontained two engineered catalysts which were placed adjacent to eachchannel wall.

The exothermic and endothermic reaction channel flows are in the samedirection (co-flow), although in this design, the reformer flow entersand leaves in a direction perpendicular to the direction of flow duringreaction to accommodate manifolding connections on a different face ofthe device than the combustion flow manifolding, just as was done forthe device of example 3. Note that the catalytic combustion was mostlikely accompanied by some homogeneous combustion in the flow-by gap.During testing, the device was insulated with ceramic fiber insulation.Two replicates of the reactor were fabricated and tested, achieving acombined time on stream of over 300 hours.

Both endothermic (combustion) and exothermic (SMR) reactions werecatalyzed by engineered catalyst inserts 134, 14 measuring 0.011″ by0.7″ by 1.0″ (0.028 cm×1.8 cm×2.5 cm) held against the channel wallswith 0.050″ (1.3 cm) wide strips of Inconel 625 metal inserted in theflow-by gap along each edge and down the center. [134 and 136 of FIG. 19b are the SMR catalyst and spacers, respectively. 14 is the combustioncatalyst. There are equivalent spacers between sets of combustioncatalysts (not shown)] The flow-by gaps in each combustion channel wereabout 0.016″ (0.041 cm) high, while the flow-by gaps in the SMR channelswere about 0.013″ (0.033 cm) high and 0.009″ (0.02 cm) high for the fulland half capacity channels respectively. The outermost SMR channels weresized to have half the flow of the inner channels since the outerchannels received only half of the heat. The thickness of theseoutermost channels were 0.018 inches and they contained an 0.011 inchengineered catalyst. After catalyst insertion, 8 and 11 (FIG. 19 b),were press-fit along each side of each SMR channel to direct the flowinto and out of the header and footer areas and prevent bypass of thecatalyst outside of the flow-by gap. Each outermost channel had catalystonly against the innermost wall and was designed to admit roughly halfthe flow going through a full channel for a given pressure drop.

Gaseous reactants were metered from pressurized cylinders via Brooksmass flow controllers, while water was delivered via high pressureliquid (HPLC) pump. An external combustor and heat exchange network wasused to continuously vaporize the feed water, preheat the steamreforming reactants to ˜800-845° C., and to preheat the combustion airto ˜500-650° C. Device body temperatures were measured in thermowells(drilled into selected webs between channels) using 0.032″ type Kthermocouples and recorded by Labview with a data logging system. Inletand outlet gas pressures were measured for the combustion and reformingstreams via calibrated pressure transducers. Gas temperatures were alsomeasured via thermocouple at the combustion gas inlet and outlet and atthe reformer gas inlet and outlet. The product gases were cooled inwater-cooled condensers and the dry exit volumetric flow was measured ina dry test meter. Dry product gas compositions (H₂, CO, CO₂, CH₄, O₂,and N₂) were measured using an online MTI GC. SMR performance wasevaluated using a 2:1 molar steam to carbon ratio. Gaseous hourly spacevelocity (GHSV) was calculated based on the entire channel volume withinwhich flow was exposed to catalyst (including the catalyst, spacerstrips, and flow-by volume) based on volumetric flows defined at 0° C.and 1 atm. Heat flux is calculated by determining the amount of heattransferred into the endothermic steam reforming reaction. For a knownmolar flowrate of reforming reactant, a known amount of conversion ismeasured through a GC analysis of the effluent composition and theoutlet product flowrate. From the total number of moles converted, thetotal heat required is calculated. Heat flux values were calculated asan average over the entire wall area in contact with SMR catalyst.Calculated average area heat flux values were based on only the requiredSMR reaction heat duty for the measured SMR conversion and selectivity.Calculated volumetric heat flux values were based upon the entire volumecontaining both the reforming and combustion catalysts and respectivegaps, including walls between channels, but not including any perimetermetal. This volume is inclusive of the entire volume through which heatis transferred between the two fluids.

Catalysts were reduced during the first day of testing at 125 C for 1hour. Startup of the device ICR was accomplished in several steps. Firstan inert flow (nitrogen, 1-5 SLPM) was initiated in both the combustionand reforming channels. Next an external combustor was ignited toprovide a hot flow of combustion products sufficient to transfer heat inthis external heat exchanger to vaporize water required for thereforming reaction and to supply required reactant preheat. Then thenitrogen flow to the combustor was shut off and air flow was initiated.Hydrogen was then introduced to the combustion channels as about 3% ofthe total flow, causing catalytic ignition of the combustion reactants.Next the hydrogen content was slowly ramped up to bring the reactortemperature to about 700° C. Then a small hydrogen flow was introducedinto the reforming channels to prevent catalyst oxidation by steam, andwater was fed through the vaporizer, thus providing steam to thereforming channels. Finally, the reformer nitrogen was turned off,methane was introduced into the reformer, and the reformer hydrogenturned off. Once the reforming reaction began, additional hydrogen wasfed to the combustion channels to maintain the reactor temperaturebetween 750 and 850° C. Changes in process flows made duringsteady-state operation were performed in a balanced fashion (changingboth the reforming and combustion flows at the same time) to preventlarge transient deviations of the device ICR temperature from thedesired operating temperature. Combustion was performed in FDR withexcess ail values as low as 13%.

This device was operated near atmospheric pressure on both the reformingand combustion reaction. The typical operating pressure oil thereforming side was 10 psig. The typical operating pressure on thecombustion side was 10 psig. Air and fuel inlet pressures were similar.

Results and Discussion

Results of tests using the high heat-flux integrated combustionmicrochannel reactor are shown in Table 5 and FIGS. 21-24. Methaneconversions approach equilibrium conversion and selectivity to carbonmonoxide and hydrogen even at very high space velocities (up to 10⁶hr⁻¹, see FIG. 21). TABLE 5 Summary of best results of tests using ICRdevice of Ex. 4 Test 1 Test 2 Test 3 Test 4 Test 5 Test 6 Test 7 airinlet temperature (° C.) 619 630 638 635 615 644 661 air inlet pressure(psig) 9.3 5.0 5.3 5.0 9.8 8.4 15.8 air inlet pressure (Pa/10⁵) 1.7 1.41.4 1.4 1.7 1.6 2.1 fuel inlet temperature (° C.) 363 363 382 483 422348 343 H2 flowrate (SLPM) 5.30 4.7 4 2.4 3.6 6.3 9.5 Air flowrate(SLPM) 29.5 20 11.4 12 20 17 30 % excess air 134% 79% 20% 110% 133% 13%33% CR contact time (msec) 3.8 5.3 8.5 9.1 5.5 5.6 3.3 CR GHSV (perhour) 958027 679979 423954 396425 649696 641438 1087416 air pressuredrop (psi) 4.3 3.7 2.0 2.4 3.2 3.3 4.6 air pressure drop (Pa/10⁵) 1.31.3 1.2 1.2 1.2 1.2 1.3 SR inlet temperature (° C.) 848 839 837 809 820843 824 SR inlet pressure (psig) 6.0 6.0 5.8 2.1 3.6 12.6 19.1 SR CH4flowrate (SLPM) 3.82 3.82 3.82 1.52 2.46 6.88 11.48 SR steam flowrate(SLPM) 7.73 7.73 7.73 3.12 4.99 13.84 23.06 Molar Steam to Carbon Ratio2.02 2.02 2.02 2.05 2.03 2.01 2.01 SR contact time (msec) 10.7 10.7 10.726.7 16.6 6.0 3.6 SR GHSV (per hour) 335612 335612 335612 134737 216393602044 1003795 Average reactor temp. (° C.) 834 850 840 854 838 854 858CH4 Conversion (GC Basis) (%) 94.9 94.6 92.1 98.1 96.0 83.7 72.6Selectivity: CO (%) 78.8 77.6 75.6 83.5 83.7 73.5 72.2 SR pressure drop(psi) 2.1 2.1 2.1 0.9 1.3 2.8 3.9 SR pressure drop (Pa/10⁵) 1.2 1.2 <1.21.1 1.1 1.2 1.3 average heat flux (W/cm{circumflex over ( )}2) 12.9 12.812.4 5.4 8.5 20.2 29.2 volumetric heat flux (W/cm{circumflex over ( )}3)52.5 52.1 50.3 21.9 34.7 81.9 118.4

NO_(x) measurements were made of the dry combustion effluent streamwhile testing the device of example 4 at the conditions shown under Test2 in Table 5. The concentration of NO_(x) measured at this condition(4.7 SLPM H₂, 20 SLPM air, 850° C. body temperature) in the dry effluentwas 10-12 ppm. This compares to NO_(x) levels exceeding 100 ppm inconventional methane steam reformers. This measurement is called herein“the standard NO_(x) test measurement.”

Results from tests in which the amount of excess air on the combustionside was varied show that the high heat-flux integrated combustionreactor is able to function effectively with 20% excess air (see FIG.22). This means that the combustion air requirement and correspondingair compressor costs and air recuperation duty can be greatly reducedrelative to traditional catalytic combustors.

The volumetric heat flux (i.e., the power transferred to the endothermicreaction per unit volume of reactor) reflects the compactness or degreeof intensification of the process. The high volumetric heat flux valuesshown in FIG. 24 demonstrates the ability of this device to drive heat(circles, left axis) into the endothermic reaction with a greatlyreduced reactor (and catalyst) volume and minimal pressure drop(triangles, right axis). The highest volumetric heat flux observed inthe high heat-flux integrated combustion reactor (FIG. 24) is 200-1000times higher than those typically seen in conventional reformers. Thegreatly reduced reactor volume suggests potential for substantialsavings in reactor materials and catalyst costs per unit volume ofsyngas produced.

Conclusions

An integrated-combustion microchannel reactor has been designed,fabricated and tested which can achieve high methane conversions at heatfluxes as high as 29 W/cm² and volumetric heat fluxes of 118 W/cm³ withpressure drops of less than 4 psi and 5 psi in the reformer andcombustor sides, respectively, per a 1 inch reactor flow length. Nearequilibrium methane conversions (73-98%) were achieved in thismicrochannel device for very high gaseous hourly space velocities(1.3×10⁵ to 1.0×10⁶) at 2:1 steam:C, ˜1 atm, and 850° C. average walltemperature. The high heat-flux microchannel reactor was shown to beable to operate with excess combustion air as low as 20%.

Example 5

An integrated-combustion microchannel reactor with distributed fuelinjection and a cross-current reforming and combustion flow orientationwas designed, fabricated, and demonstrated. This example describes thedesign and testing of this microchannel reactor. Combustion of hydrogenin air was used as the exothermic reaction, while steam methanereforming (SMR) with a steam to carbon ratio of 3:1 was used as theendothermic reaction. The device included a single SMR flow-by channelwith three combustion flow-by channels flowing in a cross-currentorientation relative to the SMR channel. SMR reactants flowed in a 0.13mm gap between the wall and a layer of porous engineered catalyst whichwas in intimate thermal contact with the solid channel wall or heattransfer surface. Combustion air flowed in three parallel 2.5 mmcylindrical channels, each with combustion catalyst coated on the wall.At three points distributed evenly along the flow length of thecombustion channels hydrogen fuel was injected into the combustionchannel. Each combustion channel also contained a static mixer made from0.5 mm notched and twisted inconel sheet material. The test results showthat this cross-current integrated combustion device can obtain highheat fluxes (˜15 W/cm²). High conversions were also possible at veryhigh space velocities.

Experimental Method

A cross-current microchannel device with integrated combustion wasdesigned, fabricated, and tested. The device body (see FIG. 25) wasfabricated from Inconel 625 by using wire EDM and conventional machiningto form both an SMR plate 202 (˜0.6 cm thick) and a combustion plate 200(1.1 cm thick), each roughly 2.3″ by 1.2″ (5.8 cm×3.0 cm). EngineeredSMR catalyst was then placed inside the SMR channel between the platesand the two pieces were welded together to seal the device. Four raisedregions, each 0.05″ (1.3 cm) wide by 0.45″ (1.1 cm) long weredistributed evenly across the SMR channel to hold the catalyst againstthe wall (heat transfer surface). The SMR catalyst measured 1.5 inch by0.45 inch by 0.011 inch thickness and were placed within a slot of equaldimension except the thickness of the slot was 0.016 inch. Externalcover plates containing the inlet and outlet gas manifolds (i.e. headersand footers) were independently machined and welded onto the body,except for the combustion fuel inlets, which were fed through threeseparate inlet tubes. The SMR header was designed to distribute the flowevenly across the 1.5″ (3.8 cm) catalyst width before flowing into the0.005″ (0.013 cm) flow-by gap. Combustion air flowed in three 0.1″ (0.3cm) diameter to cylindrical flow-by channels in a cross-currentorientation relative to the SMR channel. Combustion air was mixed withhydrogen fuel, injected from the wall opposite the flowed in threeparallel 2.5 mm cylindrical channels, each with combustion catalystcoated on the wall. At three points distributed every 0.5″ (0.13 cm)along the flow length of each combustion channel hydrogen fuel wasinjected into the combustion channel through 0.012″ (0.0030 cm) holes. Astatic mixer made from 0.5 mm notched and twisted inconel sheet materialwas inserted inside each combustion channel which enhanced mixing of theair and fuel within each channel. Combustion catalyst was coated on thestatic mixer and the wall of the combustion channels to promotecatalytic combustion of the fuel/air mixture. During testing, the devicewas insulated with ceramic fiber insulation.

Gaseous reactants were metered from pressurized cylinders via Brooksmass flow controllers, while water was delivered via high pressureliquid (HPLC) pump. An external combustor and heat exchange network wasused to continuously vaporize the feed water, preheat the steamreforming reactants combustion air to the values shown in the Tablebelow. Device body temperatures were measured in thermowells (drilledinto the outside of the combustion plate) using 0.032″ type Kthermocouples and recorded by Labview with a data logging system. Inletand outlet gas pressures were measured for the combustion and reformingstreams via calibrated pressure transducers. Gas temperatures were alsomeasured via thermocouple at the combustion gas inlet and outlet and atthe reformer gas inlet and outlet. The product gases were cooled inwater-cooled condensers and the dry exit volumetric flow was measured ina dry test meter. Dry product gas compositions (H₂, CO, CO₂, CH₄, O₂,and N₂) were measured using an online MTI GC. SMR performance wasevaluated using a 2:1 molar steam to carbon ratio. Gaseous hourly spacevelocity (GHSV) was calculated based on the entire channel volume withinwhich flow was exposed to catalyst (including the catalyst, spacerstrips, and flow-by volume) based on volumetric flows defined at 0° C.and 1 atm. Average heat flux was calculated based on the entire wallarea in contact with SMR catalyst and adjacent to plane beneath thecombustion cylinders, therefore defining the plane through which heat istransferred between fluids.

Catalysts were reduced during the first day of testing at 125° C. for 1hour. Startup of the small superchannel was accomplished in severalsteps. First nitrogen flow was started in the SMR side and air flow wasinitiated in the combustion side. Hydrogen was then introduced to thecombustion channels as about 3% of the total flow, causing catalyticignition of the combustion reactants. Next the hydrogen content wasslowly ramped up to bring the reactor temperature to about 400° C. Thena small hydrogen flow was introduced into the reforming channels toprevent catalyst oxidation by steam, and water was fed through thevaporizer, thus providing steam to the reforming channels. Finally, thereformer nitrogen was turned off, methane was introduced into thereformer, and the reformer hydrogen turned off. Once the reformingreaction began, additional hydrogen was fed to the combustion channelsto maintain the reactor body temperature between 725 and ˜875 ° C.Changes in process flows made during steady-state operation wereperformed in a balanced fashion (changing both the reforming andcombustion flows at the same time) to prevent large transient deviationsof the temperature from the desired operating temperature. To aid in theanalysis of hydrogen conversion, a quenching steam at about 400° C. wassometimes injected directly into the combustion footer. This was notvery effective at quenching the combustion in the footer of the device,and significantly increased the thermal losses when used.

Results and Discussion

Results of tests using the cross-flow integrated combustion microchannelreactor are shown in the Table below and FIG. 26. Important featuresdemonstrated include: catalyst applied directly to wall; static mixerinserted into reaction chamber; distributed fuel; and cross-floworientation of combustion fluids and reformate. TABLE Summary of bestresults of tests using ICR device of Ex. 5 CR inlet temperature (° C.)799 799 776 CR inlet pressure (psig) 6.43 10.94 4.83 CR inlet pressure(Pa/10⁵) 1.5 1.8 1.3 CR fuel inlet temperature (° C.) 25 25 25 H2 flowrow 1 (SLPM) 0.36 0.36 0.30 H2 flow row 2 (SLPM) 0.36 0.36 0.30 H2 flowrow 3 (SLPM) 0.36 0.36 0.30 air flow rate (SLPM) 9.0 14.4 5.2 400 C.quench steam flow (SLPM) 21.1 21.1 14.9 % excess air  250%  460%  143%CR contact time (msec) 1.15 0.75 1.90 SR GHSV (per hour) 3132777 48110511895827 air pressure drop (psi) 1.1 1.3 0.5 air pressure drop (Pa/10⁵)1.1 1.1 1.0 H2 Conversion (GC basis) (%) 99.9% 99.9% 99.9% SR inlettemperature (° C.) 817 807 825 SR inlet pressure (psig) 14.3 13.9 14.7SR inlet pressure (Pa/10⁵) 2.0 2.0 2.0 SR CH4 flow (SLPM) 0.53 0.53 0.53SR steam flow (SLPM) 1.59 1.59 1.59 Molar Steam to Carbon Ratio 3.0 3.03.0 SR contact time (msec) 5.0 5.0 5.0 SR GHSV (per hour) 720670 720670720670 Average IC body temperature (° C.) 856 844 856 Apparent SMRtemperature (° C.) 766 753 837 CH4 Conversion (GC Basis) (%) 73.8 73.381.1 Selectivity: CO (%) 62.2 61.2 67.2 SMR pressure drop (psi) 10.810.3 11.3 SMR pressure drop (Pa/105) 1.8 1.7 1.8 SMR heat flux(W/cm{circumflex over ( )}2): 13.6 13.9 16.5 Heat load per unit volume(reactor zone) (W/cm3) 34.0 34.8 41.3For all tests the conversion of hydrogen in the combustion reactionchamber was 100%

Example 6

Design and Operation

An integrated reforming and combustion reactor was evaluated in across-flow orientation. A single SMR slot was adjacent to threecombustion cylinders.

The SMR engineered catalyst was placed against the wall shared with thecombustion channels, which provides the energy supplying the endothermicreforming reaction. The SMR channel gap was 0.015 inches and an SMRporous engineered catalyst of 0.012 inches was inserted leaving a gap of0.003 inches. The length of this reforming channel was 0.45 inches andthe width was 0.5 inches. The combustion section was made up of three0.100″ (0.254 cm) diameter channels in a plane parallel with the SMRchannels, but with its flow path aligned 90° to the flow of the SMRchannel, in cross-flow fashion. There was a 0.05 inch web of metalseparating the reforming channel and the top of the combustioncylindrical channels. The flow length of the combustion cylinders was0.5 inches which matched with the 0.5 inch width of the reformingchannel for this cross-flow orientation. The combustion catalyst wasmade from a 0.010″ (0.025 cm) thick FeCrAlY felt wrapped around a 0.060″(0.15 cm) wide and 0.020″ (0.051 cm) thick static mixer insert. Themixer served two functions: (1) pressing the felt against the chamberwalls and (2) mixing the fuel as the air passed through the innerdiameter of the felt.

The static mixer was made in the following manner:

-   -   1. The 0.060″ wide side was cut into two 0.250″ length sections        separated by a 0.020″ width that extended 0.040″ in length.    -   2. The first 0.250″ length section was twisted 90° from its        beginning orientation    -   3. The next (and last) 0.250″ length section's beginning was        turned 90° to tile end of the first 0.250″ length section.    -   4. The end of the second 0.250″ long length was twisted 90° in        the opposite direction of the first 0.250″ length section.        The fuel is fed into each combustion channel at the inlet of the        cylindrical channel from the bottom of the device, or on the        side opposite the plane that separates the reforming channel.        Combustion fuel enters 0.030″ (0.076 cm) ahead of the        catalyst/static mixer insert placed within the combustion        cylindrical channel, each channel with it own fuel port.

The overall dimensions of the block device in the plane of the channelsare 1.65″ across in the SMR channel direction and 1.700″ in thecombustion channel direction, centered over the 0.450″ by 0.500″integrated combustion reforming core where the reactions occur and heatis transferred. The volume through which all heat is transferred andused for volumetric heat flux calculations is 0.45 inches by 0.5 inchesby the sum of 0.1 inch (combustion diameter) and 0.05 inch (metal webseparating channels) and 0.015 inches (full reforming channel) for atotal of 0.037 in3 or 0.6 cm3. The area or plane through which all heatis transferred is 0.45 inches by 0.5 inches or 1.45 cm². The processheader was specially made to have a smooth transition from the innerdiameter of a 0.180″ inner diameter inlet tube to the 0.500″ wide by0.014″ tall slit to avoid coking and allow additional preheating fromdevice losses. A footer of the same design leads to the outlet tube of0.180″ inner diameter.

This steam methane reforming microchannel reactor supported a pressuredifferential of 11 atm absolute at 850 C between reforming andcombustion channels. The metal web separating channels was 1.27 mm or1270 microns. The device was operated for more than 13 days or more than300 hours without any change in structural integrity. No leaks werefound between streams. The high pressure differential between streamswas supported over a thin metal web by the use of microchannel,especially circular microchannels, beneath a high pressure reformingslot.

Performance

The device was demonstrated for the reforming reaction at 3:1 molarsteam to carbon ratio, 12 atmospheres absolute pressure and 850° C., and5 milliseconds contact time. The device produced equilibrium conversion,as seen in FIG. 27. The methane conversion and CO selectivity decreasedcontinuously over time for the next five and a half days. The conditionswere changed to 2:1 molar steam to carbon, with the temperature,pressure and contact time held constant, and the loss of activitycontinued. This device did show substantial heat fluxes and averagevolumetric heat fluxes, as shown in Table 5. The first two settingsrefer to the 3:1 steam to carbon ratio at 5.0 millisecond contact times,with setting # 1 referring to the beginning of the run in FIG. 27 andsetting #2 referring to the end of the run. Similarly, settings #3 and#4 in the Table below refer to the beginning and end performance of the2:1 steam to carbon molar ratio at 5 milliseconds contact time, alsoillustrated in FIG. 27. All data reflect an SMR reactant inlettemperature of 830 to 840 degrees Celsius. TABLE 5 Total Average MethaneWater Inlet Process CO Total heat Average volumetric flow rate flowratepressure dP Methane Selectivity load heat flux heat flux Setting #(SLPM) (ccm) (Psig) (psid) Conversion % (%) (W) (W/cm²) (W/cm³) 1 0.1530.37 174.3 2.75 91 63.83 22.2 15.3 43.1 2 0.153 0.37 174.5 2.65 75.458.1 18.2 12.6 35.3 3 0.206 0.33 174.6 2.73 70.1 68.9 23.2 16.0 45.0 40.206 0.33 173.7 2.74 56.5 64.4 18.6 12.8 36.0

The corresponding settings on the combustion side to those in Table 5are shown in Table 6. They reflect a constant air flow rate of 5.4 SLPMfor three channels, and a fairly constant hydrogen flow rate near 0.5SLPM. TABLE 6 Adiabatic Hydrogen Air Air inlet Excess Flame flow rateflow rate temp Air Temp Setting # (SLPM) (SLPM) (Celsius) (%) (Celsius)1 0.506 5.4 817 348 1450 2 0.508 5.4 771 347 1411 3 0.514 5.4 768 3421412 4 0.514 5.4 788 341 1432For all tests, the conversion of hydrogen in the combustion reactionchamber was 100%.

Heat Flux Measurement Test

Operate the device for a methane steam reforming reaction at 850 C, 1.70bar (10 psig), 3:1 steam to carbon ratio, and a contact time of 100 ms.Contact time is defined as the total reaction volume divided by thetotal volumetric inlet flowrate of reactants at standard temperature andpressure (STP: 273K and 1 atm absolute). The total reaction volume isinclusive of the steam reforming catalyst and the reaction channelvolume that contains the catalyst.

For example, if the cumulative volumetric sum of reaction chambersinclusive of reforming catalysts is 1 cubic centimeter, then the inlettotal flowrate of reactants would be 0.6 standard liters per minute. Theinlet flowrate of methane would be 0.15 standard liters per minute andthe inlet flowrate of steam would be calculated to be 0.45 liters perminute at standard temperature and pressure. For this example, the inletmolar flowrate of methane would be roughly 0.00045 moles per second.These numbers scale linearly with the total reaction chamber volume. A 2cubic centimeter reaction chamber volume would require 0.0009 moles persecond.

Methane conversion is determined by measuring the outlet productcomposition and the outlet flowrate of methane reforming reactionproducts and then calculating based on the following formula.Conversion %=100×(moles methane in−moles methane out)/(moles methane in)Moles methane in=inlet flowrate of methane at STP/(22.4 L/mol)Moles methane out=[outlet flowrate of total product dry gas/(22.4L/mol)]×% methane in dry gas GC analysis

Dry gas is defined as the product gas stream flowrate after condensingthe unreacted water or other condensable fluids.Selectivity to CO %=100×(moles of CO/(moles of CO2+moles of CO+moles ofC(s) if present))Selectivity to CO2% 32 100×(moles of CO2/(moles of CO2+moles of CO+molesof C(s) if present))Heat load=(Conversion %/100)×Moles methane in×(Heat of reaction of steamreforming to carbon monoxide at 850 C (226800 J/mol)×selectivity to CO%+Heat of reaction of steam reforming of methane to carbon dioxide at850 (193200 J/mol)×selectivity to CO2%)/100, units of WattsHeat flux=Heat load/reactor core volume, units of Watts/cm3

Where the reactor core volume includes all reaction chambers orchannels, all associate combustion chambers or channels, and allseparating metal webs through which heat transfers between fluids. Inshort, this volume includes the total volume through which heattransfers for the methane steam reforming reaction. This volume does notinclude perimeter metal, manifold volume, or other associated packagingthat is dependent on individual device geometries.

The following conditions must be met for the combustion reaction thatsupplies heat for the heat flux measurement test:

-   -   1. The gas phase fuel that must be used is hydrogen.    -   2. The total air flow rate is sized such that a mixture of the        hydrogen and air flow rates into the reactor reaches an excess        air percentage of 80%. The excess air is defined as the total        molar flow rate of oxygen in the combination of hydrogen and air        divided by the molar flow rate of oxygen needed to fully oxidize        the hydrogen at its molar fuel flown rate. As one mole of oxygen        can fully oxidize two moles of hydrogen, 80% excess air        corresponds to a 4.28:1 molar ratio of air to hydrogen. Air is        taken as 21% mole percent oxygen, balance nitrogen.    -   3. The hydrogen and air enter the combustion reactor at 900° C.    -   4. The air and hydrogen are to be mixed either in a manifold        that is directly upstream of the combustion reactor or in the        reactor itself.    -   5. The standard volumetric flow rates for hydrogen though the        combustion reactor per 0.15 SLPM of methane flow rate through        the methane steam reforming reactor is a minimum of 0.140 SLPM        and a maximum of 0.204 SLPM.    -   6. The corresponding minimum and maximum air flow rates through        the combustion reactor, based upon the 80% excess air condition,        per 0.15 SLPM of methane flow rate through the methane steam        reforming reactor is 0.600 SLPM and 0.875 SLPM, respectively.    -   7. The inlet pressures of the hydrogen and air streams should be        no greater than 2.38 bar (20 psig).        Pressure Test—High Temperature Test for ICR

In preferred embodiments, any of the devices described herein arecapable of withstanding internal pressure differences. For example, somepreferred embodiments meet the requirements of the following pressuretest. For a microchannel unit operation device with at least onecritical channel dimension less than about 2 mm, operate with at leasttwo inlet fluid streams. The first fluid stream must be at 850 C and 180psig. The second fluid stream must be at 800 C and 10 psig. Any flowrate may be used. Operate the device with gas flow to both streams for300 hours. After 300 hours operation, pressurize each fluid flow line to50 psig and hold for 2 hours. The pressure must remain constantindicating minimal leak paths to the environment. Then, pressurize thesecond fluid flow line to 50 psig, leaving the first fluid flow lineopen to atmosphere, and hold for 2 hours. The pressure must remainconstant indicating minimal internal leak paths. A minimal leak path isdefined as a leak rate of less than 10⁻⁶ standard cubic centimeters persecond of helium when helium is used as the fluid for the final leaktest.

The invention also includes methods of conducting unit operations in thedevice having the pressure resistance characteristic described above.

1. An integrated reactor, comprising: a first reaction chamber having awidth of 2 mm or less, wherein there is an open path through the firstreaction chamber, wherein the first reaction chamber has an internalvolume comprising 5 to 95 vol. % of porous catalyst and 5 to 95 vol. %of open space; and a second reaction chamber having a width of 2 mm orless, wherein there is an open path through the second reaction chamber,wherein the second reaction chamber has an internal volume comprising acatalyst and at least 5 vol. % of open space; and a reaction chamberwall separating the first chamber and the second chamber; and whereinthe integrated reactor possesses a heat flux characteristic of at least1 W/cc as measured according to the Heat Flux Measurement Test.
 2. Thereactor of claim 1 wherein the heat flux characteristic is obtained witha pressure drop of less than 12,500 Pa/cm.
 3. The reactor of claim 1wherein the second reaction chamber has an internal volume comprising 5to 95 vol. % of porous catalyst and 5 to 95 vol. % of open space.
 4. Thereactor of claim 1 wherein the second reaction chamber comprisesreaction chamber walls and a catalyst wash coated onto at least aportion of said reaction chamber walls.
 5. An integrated reactor,comprising: a first reaction chamber having a width of 2 mm or less,wherein there is an open path through the first reaction chamber,wherein the first reaction chamber has an internal volume comprising 5to 95 vol. % of porous catalyst and 5 to 95 vol. % of open space; and asecond reaction chamber having a width of 2 mm or less, wherein there isan open path through the second reaction chamber, wherein the secondreaction chamber has an internal volume comprising a catalyst and atleast 5 vol. % of open space; and a reaction chamber wall separating thefirst chamber and the second chamber; and wherein the integrated reactorpossesses a NOx output characteristic of less than 100 ppm as measuredaccording to the Standard NOx Test Measurement.
 6. A method of making anintegrated reactor, comprising: providing a single block of thermallyconductive material; forming at least one first microchannel; forming atleast one second microchannel; placing at least one catalyst capable ofcatalyzing an exothermic reaction in the at least one firstmicrochannel; placing at least one catalyst capable of catalyzing anendothermic reaction in the at least one second microchannel; whereinthe first microchannel and second microchannel are separated by lessthan 1 cm. 7-17. (canceled)